Solvent extraction process for separating cobalt from nickel in aqueous solution

ABSTRACT

A process for separating Co from Ni in an aqueous solution comprises subjecting the solution to extraction and using kinetic differences between Ni and Co in the extraction for achieving at least a partial separation of Co from Ni. This is effected by controlling the duration of the extraction so that a major portion of Co and a minor portion of Ni is extracted from the solution to produce a loaded extractant, enriched in Co and depleted in Ni compared to the feed solution, and a Co-depleted raffinate containing Ni. In a further embodiment, the invention utilizes kinetic differences between Ni and Co during striping for effecting separation of Ni and Co. The loaded extractant can be subjected to a bulk stripping or a selective stripping operation to obtain Co and Ni solutions from which Ni and Co can be recovered. The process may be incorporated in a hydrometallurgical process for the extraction of Ni and/or Co from an ore or concentrate containing Ni and Co.

CROSS REFERENCE TO RELATED APPLICATION

This application is a continuation of U.S. patent application Ser. No.12/784,037 filed May 20, 2010 entitled SOLVENT EXTRACTION PROCESS FORSEPARATING COBALT FROM NICKEL IN AQUEOUS SOLUTION, the disclosure ofwhich is hereby incorporated by reference in its entirety.

FIELD OF THE INVENTION

The present invention relates to a process for separating cobalt fromnickel in an aqueous solution.

BACKGROUND OF THE INVENTION Properties of Cobalt and Nickel in AqueousSolution

Cobalt and nickel have very similar chemical properties. In aqueoussolution both elements are normally found as divalent cations (e.g. Co⁺⁺or Ni⁺⁺) in acid solution, with similar solubilities.

Both are soluble as the respective sulphates, chlorides, or nitrates inacid solution, but are largely insoluble in alkaline conditions unlesschelating agents are present. (Both cations are chelated strongly byammonia, for example). Their sulphide compounds have similar properties,(the K_(SP) values are similar leading to a similar pH of formation withS²⁻ ions), as do their respective carbonates.

This phenomenon presents an ongoing problem for the extractivemetallurgist, as Co and Ni are invariably found together in naturallyoccurring ores, but must ultimately be separated to make maximum use ofeach metal. Fortunately however, there are exceptions to this pattern ofsimilar behavior, which can be exploited, and will be discussed below.

Ratio of Cobalt to Nickel Occurring Naturally in Nickel Ores

Ni and Co commonly occur together in nature as sulphide ore deposits,and the ratio of Co:Ni in Ni ores is surprisingly constant in the rangeof about 1:15 up to 1:30. This is true at least in unaltered sulphideores; although this ratio can be quite different if weathering of theores has occurred, (over many millennia),

A general review of nickel metallurgy is to be found in the excellentbook “The Winning of Nickel”, by Joseph P. Boldt Jr., and Paul QueneauSr., published by Methuen and Co. 1967. A more recent review oflaterites in particular is found in “The Past and Future of NickelLaterites,” by Ashok Dalvi, Gordon Bacon, and Robert Osborne in: PDAC2004 International Convention, Trade Show and Investors' Exchange, (Mar.7-10, 2004).

Weathering of surface Ni deposits is common especially in tropicalcountries, and such deposits are usually referred to as Ni laterites;due to slight differences in chemistry, this weathering frequentlyresults in a partial separation of Ni and Co over a vertical horizon,compared to the original sulphide ore. Some concentration of Co into theso-called limonite layer often is the result, so Co:Ni ratios inlaterites considered for hydrometallurgical processing varysubstantially from the usual ratio in sulphides, e.g. a ratio of 1:10 oreven lower may be found in limonite.

Co Separation from Nickel in a Leach Solution . . . Overview

The distribution of Co and Ni in most leaching processes is verysimilar. Thus leaching of Ni—Co ores or concentrates usually results ina mixed solution of Co and Ni, as well as other materials.

However, Ni and Co have to be separated eventually to obtain maximum useand payment for each element, as their end-use is significantlydifferent, e.g. different metal alloys. With only a few notableexceptions, Ni is much more abundant in ores or concentrates than Co.Thus in a mixed Ni—Co solution, the problem of separation can be moreaccurately stated as separating Co (as an impurity, albeit of value)from a Ni solution. It also happens that Co has a few specific chemicalproperties that allow for its selective extraction from a mixture ofelements in solution, whereas Ni in general does not have suchproperties.

A variety of processes have been (and some still are) used commerciallyto achieve this objective, but all of them have significant costs, andit is the objective of the present invention to provide a more efficientand more cost-effective method.

Also, because of the greater value of Co, this need to purify the Nisolution of Co should not obscure the secondary need to also recover Coitself in an economic manner, which is part of the present invention.

A number of processes have been used commercially for Co separation fromNi, for example:

Precipitation of Co from Ni—Co solution as Co(OH)₃

This is an old process, (see the above-referenced book by Boldt andQueneau), one of the first known methods, and still used commercially.With strong oxidants, a Co⁺⁺ solution can readily be oxidized to Co⁺⁺⁺,which is essentially insoluble in dilute acid solution (say pH 2-6). Cois thus precipitated as Co(OH)₃ whilst the Ni stays largely in solution.Oxidants used for this purpose include Cl₂ and ozone. Electro-oxidationcan also be used. However, the process is costly and inefficient, due tosignificant co-oxidation of Ni⁺⁺ to a similar product, and has generallyfallen out of favour.

Selective Solvent Extraction (SX) of Co from Ni—Co Solution

This approach has been the subject of many investigations, some of whichhave been commercialized, and some of these are listed in the referencedarticle, “Cobalt-Nickel Separation in Hydrometallurgy: a Review,” byDouglas S. Flett, in: Chemistry for Sustainable Development 12 (2004),pages 81-91. There are some organic extractants which will selectivelyextract Co with respect to Ni. Primarily these are one of two types:

i) Ternary and quaternary amines (Alamine™ 336 for example), which canextract some metal chloride complexes (e.g. CoCl₄ ²⁻) from a strongchloride aqueous environment. Unusually, Ni doesn't form such chloridecomplexes, so a good separation of Ni from other elements can sometimesbe achieved. However, the requirement of the strong chlorideconcentration (several molar) severely limits the applicability of theprocess, and in reality omits it from serious consideration for atypical leach liquor.

ii) Phosphinic acids (e.g. bis 2,4,4-trimethylpentyl phosphinic acid,sold commercially as Cyanex™ 272), which will extract Co selectivelyover Ni, without the need for a chelating agent as in the amines.Although this extractant works well in pure solutions, unfortunately italso extracts many other metals commonly found in leach solutions, suchas Mg and Mn, which limits its usefulness. This limitation isillustrated in the first two Examples described below. Although Mg canbe scrubbed off the loaded organic stream by Ni/Co (with difficulty), byusing a large number of mixer/settlers in counter-current mode, as wasdone at the Bulong plant (described by Donegan (“Direct SolventExtraction of Nickel at Bulong Operations,” by S. Donegan, in: MineralsEngineering 19 (2006), pages 1234-1245), Mn cannot be scrubbed off. Theonly remedy is to co-extract all the Mn along with the Co, and pay forthe cost of the Mn extraction/stripping, particularly the cost of theammonia used for neutralization of the organic extractant. As a resultof this limitation, Cyanex 272 is best applied after some priorpurification, which itself is both costly and inefficient.

Selective SX of Ni from Ni—Co solution

This approach was invented and commercialized at the Queensland Nickel(QNI) plant in Yabulu, Queensland, Australia, very successfully in the80's, and is well summarized by the above-referenced Flett reviewarticle, and also by Reid and Price. (Reid, J G and Price, M J, 1993.Ammoniacal solvent extraction at Queensland Nickel: Process installationand operation, in Solvent Extraction in the Process Industries Volume 1(Proceedings of International Solvent Extraction Conference 1993) (eds:D H Logsdail and M J Slater), pp 225-231 (Elsevier Applied Science:London and New York.

The process is effective but suffers from high cost, as the majorcomponent (Ni), is being extracted away from the ‘impurity’ (Co),generally a more expensive route than the opposite. The impurities areall left with the Co.

Also the extractant used (hydroxyoxime) is prone to rapid degradation byCo II oxidation, and hence must be re-oximated on a regular basis, atconsiderable cost. Pre-oxidation of Co II to Co III is necessary tominimize this problem, but is not 100% effective, leading to continuousre-oximation of the extractant, at high cost.

Stripping of the Ni from the loaded organic can be done with eitherstrong ammonia solutions (250 g/l NH₃), or by acid, as was done at theCawse mine in Western Australia for a while, as discussed by Flett(referenced above). The former fits in well with NiCO₃ production, (bysteam stripping of NH₃), the latter with a Ni electrowinning flowsheet.

Although technically feasible, this approach is relatively expensive asnoted, and does not produce a pure Co product, as the Ni left in theraffinate produces a Ni:Co ratio in this stream of at least 1:1. AtYabulu, a separate Co refinery had to be built eventually forre-processing of the Co-rich stream, and this refinery had its owntechnical and financial challenges.

Hydrogen Reduction of Ni from Ni—Co Solution

This approach was first commercialized in about 1950 at the SherrittGordon plant in Fort Saskatchewan, Alberta, as described by by Boldt etal. The method has been considered the standard process for nickelrecovery by some designers; it has been installed in several othernickel plants since, but suffers from significant drawbacks:

-   -   Batch mode operation. The process apparently can only be        operated in this way, (instead of the usual continuous mode).        This then requires multiple units (autoclaves), with low        operating time (need to fill and discharge each batch autoclave        frequently)    -   These features lead to high capital and operating costs    -   Heightened safety and occupational health requirements    -   Need for concentrated Ni solution (50-100 g/l [Ni]), and high        solution feed temperatures, approx 200° C., again leading to        high costs.    -   Need for careful control of pH, in pH 7.0 range, with high        background levels of ammonium sulphate (200 g/l).    -   Large by-product production of ammonium sulphate crystals, at        about 7× tonnage of Ni metal production, necessitating        evaporative crystallizers, filters, dryers, bagging, storage        facilities, etc. All leading to high costs.    -   Technically complex, thus requiring high level of technical        skill, and expensive engineering input from a very limited        number of qualified suppliers    -   Further processing costs, downstream from the actual hydrogen        reduction due to need for further purification to remove trace        amounts of impurities such as S and O. High temperature        oxidation and reduction furnaces are needed for this        purification.    -   Relatively poor quality of Ni product, regarding Co content    -   Poor quality of Co product, as the raffinate contains a Ni:Co        ratio of about 1:1, leading to need for another Co refining        step, similar to the process described in previous section.        Difficulties with Separation of Cobalt from Nickel in the        Presence of Impurities

As described in the previous section, Co separation from Ni isdifficult, and specific to each situation, e.g. solution chemistry andparticularly the impurities in the Co—Ni solution. For solutions derivedfrom leaching sulphide concentrates (as at Sherritt Gordon, forexample), impurities are generally confined to other base metals such asFe, Cu and Zn, which can be removed efficiently by known purificationmethods.

However, for acidic solutions derived from leaching of laterite ores,other impurities are found, particularly Mg and Mn, and often in muchgreater concentrations relative to the Ni and Co concentrations, e.g. asmuch as 10× greater. This situation makes it near difficult to use thepreferred Cyanex 272 Co extraction method described above without someform of pre-treatment to separate Co and Ni from these impurities, oralternatively pay for the expense of co-extracting Mn.

Thus, unless Mn extraction is to be tolerated and paid for, treatment oflaterite leach liquors is usually required to choose one of two routes:

-   -   Precipitation of Ni and Co in acid solution away from impurities        as much as possible, followed by re-leaching to form a new        solution with reduced impurity content. This approach for        example was followed in the Murrin Murrin plant, which is        described by Campbell et al in U.S. Pat. No. 7,387,767, and also        in the Cawse plant, described by White in U.S. Pat. No.        6,409,979.    -   Leach in an ammoniacal alkaline environment, wherein most of        damaging impurities are largely absent. This approach is adopted        by the Caron process, for example (described in Boldt and        Queneau), which was used at the QNI plant in Yabulu mentioned        above, (now under different ownership and renamed).

The Caron process requires a pretreatment process of its own, areductive roast at high temperature. This is a pyrometallurgicalprocess, and requires high capital investment. It also has high energyrequirements and thus has high operating costs. For these reasons, it isgenerally not considered today, although a few plants built years agoare still operating.

Practically then, one is confined to the precipitation and releachoption as a pre-treatment prior to Co—Ni separation.

For laterite ores therefore, it is desirable, even necessary, that theMn/Mg be separated out from Ni/Co by first precipitating the Ni/Co fromsolution, and then releaching. This is usually done by one of twomethods:

-   -   1. H₂S precipitation of mixed Ni/Co sulphides, (which        selectively precipitates Ni/Co over Mn/Mg), followed by        filtration and then pressure oxidation of mixed sulphides        precipitate, with associated filtration steps, to produce a        Ni/Co solution suitable for efficient Cyanex 272 extraction of        Co, (as is now done at Murrin Murrin, op cit), or    -   2. Mixed hydroxide precipitation of Ni/Co with MgO (which can be        done selectively over Mn and Mg, as practiced at Cawse plant for        example, and patented by White, op cit), or with CaO, which is        not so selective with respect to Mg/Mn. At Cawse, this was        followed by re-leaching of mixed hydroxides with ammonium        carbonate solution, filtering, then steam stripping the NH₃/CO₂        from the leach liquor to precipitate Mn/Mg, (and re-adsorbing        the same NH₃/CO₂), and re-filtering.

Either of these processes for Mn/Mg rejection is expensive, especiallyin capital costs (H₂S generating plant, or stripping and absorptionplants for NH₃/CO₂). It is an objective of the present invention to beable to treat high Mn/Mg solutions containing both Ni and Co, andseparate Co from this solution, without going through either of theexisting Mn/Mg rejection alternatives sketched out above.

Unfortunately, a large amount of gypsum is formed along with the mixedNi—Co hydroxide precipitate (MHP) when slaked lime is used as theprecipitant for Ni and Co; consequently the MHP contains about 50 wt %gypsum, and only 50 wt % actual Ni and Co hydroxides, (and hence about22% Ni).

Also both Mg and Mn are partly precipitated from solution (ashydroxides) at about the same pH as Ni and Co, thus furthercontaminating the product. This is a particular problem for lateriteswhere these impurities are usually present in high concentration, e.g.in leach liquors derived from laterites by a High Pressure Acid Leach(HPAL) process.

Mn is of special interest as well because MHP is of course a mixture ofNi and Co compounds, which eventually have to be separated to makecommercial Ni and Co products. The conventional technology for thisseparation is to use solvent extraction on a Ni/Co solution, inparticular the extractant Cyanex 272. This Cyanex 272 separation is onlyfeasible if Mn and Mg are very low in the feed solution, otherwise theyinterfere with the Co extraction. Generally this limits such refining tofeed materials that are already low in Mn: Ni sulphide concentratesgenerally fall into this category, or are smelted to matte anyway, whichis an effective Mn removal step, (into the slag). For feed materialsthat have not been smelted, e.g. laterite high pressure acid leachliquors, a significant Mn presence may be a serious impediment tofurther processing.

Commercial Value of Nickel Hydroxide Products

Mixed Ni—Co hydroxide (MHP) has been produced by at least one Ni mine inthe recent past, and several announced projects have included thisintermediate in the process flowsheet. However, this product hasuncertain market value and a limited marketing history, due in part toits purity and grade. Since Ni and Co must be separated eventually, thepresence of other impurities in the MHP can be a serious impediment tosaid separation, as described in the previous section.

If slaked lime, (Ca(OH)₂), is used as the reagent, it typicallyprecipitates some Mg and Mn from solution along with Ni and Co; and ifany traces of base metal impurities, (e.g. Cu, Fe, Zn and Cd), are leftin solution from the prior purification, they are also precipitated intothe MHP. In addition, gypsum is formed of course, thus degrading the Nigrade of product by about 50%. If MgO is used instead of slaked lime,better selectivity is found in respect to Mg, and of course no gypsum isformed, so the Ni grade is much better, but most of the same impurityissues remain.

All of the above can affect on the marketing of the MHP, and hence itscommercial value. Nevertheless for projects with limited Ni/Coproduction it may be advantageous to be able to sell the MHP as anintermediate product, and thus avoid the extra cost of full refining tometal products on a small scale. It was of some importance therefore toimprove the quality of MHP, by Co separation upstream.

Therefore if the Co is separated out before the precipitation process,the resultant Ni Hydroxide Product (NHP) may have enhanced value.

U.S. Pat. No. 6,171,564 relates to a process for treatment of nickelores and concentrates to recover both Ni and Co as refined metals. It isa ‘comprehensive’ process in the sense that:

-   -   a) Both sulphides and oxides (laterites) are considered as        suitable feeds, and    -   b) The process describes a complete flowsheet going all the way        to metal product:        -   Acid leaching of solid feed to produce an acidic leach            solution, containing Ni, Co and numerous impurities        -   purification of leach solution, in several steps, including            solvent extraction        -   precipitation of Ni and Co together as an impure mixed            hydroxide by neutralizing acidic solution        -   Releaching of mixed hydroxide in (recycled) ammoniacal            solution to redissolve Ni and Co values        -   Separation of Co from Ni in said solution by solvent            extraction, followed by stripping and recovery of Co product        -   Further purification of (ammoniacal) raffinate from Co            solvent extraction by additional solvent extraction steps        -   Extraction of Ni from purified ammoniacal solution, followed            by acidic stripping of organic stream to form purified            acidic Ni solution        -   Ni recovery from purified solution by electrowinning, with            recycling of spent acid to solvent extraction

The Ni and Co content in the feed material are first leached by pressureoxidation (in the case of sulphides) or by acid pressure leaching (inthe case of laterites); then the solution is purified to removeprimarily Cu, Zn and Fe. From the purified solution, Ni and Co areprecipitated together at about pH 7-8, using slaked lime, as a mixtureof Ni and Co hydroxides, (MHP).

MHP is then re-leached in mild conditions, (ambient temperature, dilutesolids, neutral pH) with a strong ammonium sulphate solution (200 g/l)at about pH 7.0, as in Reaction (1):

Ni(OH)₂+(NH₄)₂SO₄→Ni(NH₃)₂SO₄+2H₂O  (1)

Thus this leach produces a solution of Ni and Co diammine, (Ni(NH₃)₂ ⁺⁺and Co(NH₃)₂ ⁺⁺ ions), which also contains some impurities, notably Mgand Ca.

The leaching of Ni and Co by this method is not very efficient, (about90%), due to the mild conditions selected, i.e. the neutral pH and hencevery low free ammonia content of the leach solution; almost none of theNi species in solution is present as free ammonia, NH₃. As aconsequence, significant Ni and Co are left behind in the residue, whichtherefore has to be releached to avoid unacceptable losses. However, thechoice of the neutral pH is very important to enable the subsequentsolvent extraction steps to proceed efficiently.

Co is then extracted selectively from this impure Ni/Co diamminesolution using as extractant the Cyanex 272 reagent at about this samepH 7.0, leaving Ni in the raffinate.

Since the extraction is from a diammine, (i.e. Co(NH₃)₂SO₄), noneutralization is required to maintain a constant pH, as the ammine issimultaneously converted to ammonia (NH₃) during the extraction, whichexactly balances the acid production (H+ ions) from the organic reagent,and thus produces ammonium sulphate in the overall extraction reaction(2):

Co(NH₃)₂SO₄+2RH(organic reagent)→CoR₂(organic phase)+(NH₄)₂SO₄  (2)

Thus the pH stays almost constant throughout the Co extraction, negatingthe usual need for neutralizing agent. This is an unusual and mostbeneficial feature of this solvent extraction process, as most other Coand Ni solvent extraction processes need in situ neutralization withammonia or caustic to counteract acid production, and thus maintain thesolution pH within the required range during the reaction, (or else theextraction stops prematurely). The significance of the neutralizationgoes far beyond the simple avoidance of reagent consumption; the normalbyproduct of such neutralization with ammonia or caustic is a salt suchas ammonium sulphate or sodium sulphate, which rapidly accumulates inthe raffinate stream, and must be disposed of in some fashion. This is aserious challenge, given the constraints of the system, such as metalcontamination of said salt as a potential byproduct, and is sometimesfatal to a process design.

Co extraction at this point is only about 90% of the Co contained in thediammine solution, and co-extraction of Ni and Mg is negligible, thusproviding a relatively pure Co stream (in the loaded organic), in theabsence of Mn, Fe, Cu or Zn (all of which can co-extract with Co). Coextraction is kept deliberately less than 100%, to ensure that theloaded organic (loaded organic) is fully loaded with Co, thus minimizingco-extraction of Mg and Ni, (which are less strongly extracted than Co).Even then, some scrubbing of loaded organic is required to remove thesmall amounts of Mg and Ni that are extracted. Scrub feed is derivedfrom a fraction of the (pure) cobalt strip liquor, which is in limitedsupply, since the Ni:Co ratio in the solution is typically >10:1; sominimizing of scrubbing requirements is essential.

Co is then stripped from the loaded organic in conventional fashion withdilute sulphuric acid solution to form a concentrated and pure Cosolution, (low in Mg and Ni), and then recovered from this stripsolution by conventional electrowinning (as pure metallic cathode), orby precipitation with some suitable reagent (e.g. sodium carbonate) as apure Co salt, carbonate or hydroxide, etc.

Cyanex 272 is applied as an extractant again at pH 7.0-7.5 to the Coraffinate, to recover the residual Co, (about 10% of feed Co in ammoniumsulphate solution), and also to remove any Mg and Ca from this solutionprior to Ni recovery. This is called the magnesium extraction stage forsake of reference. About 10% of feed Ni is also extracted here into theloaded organic, leaving about 90% of feed Ni in raffinate from thisoperation. Stripping of the loaded organic stream with acid produces anacidic aqueous stream which is recycled to the leach circuit forrecovery of Ni/Co values. Co in raffinate is very low, (˜1 ppm [Co]), inorder to produce high purity Ni in the next step.

Ni is extracted from Mg raffinate with LIX™ 84 extractant at about pH7.0-7.5, as in Reaction (3);

Ni(NH₃)₂SO₄+2RH(organic reagent)→NiR₂(organic phase)+(NH₄)₂SO₄  (3)

The Ni extraction is followed by acidic stripping of the loaded organic,to produce a pure Ni electrolyte, (4), and recovery of Ni metal ascathodes by conventional electrowinning, reaction (5):

NiR₂(organic phase)+H₂SO₄→2RH(organic phase)+H₂SO₄+H₂O  (4)

NiSO₄+H₂O→Ni⁰↓+H₂SO₄+½O₂  (5)

The final Ni raffinate is recycled to the original ammonium sulphateleach, completing the circuit.

It is to be noticed that the overall reaction, i.e. combining leachingof Ni hydroxide, solvent extraction, solvent stripping andelectrowinning, reaction (6), has no reagent consumed and no byproducts,other than water and oxygen:

Ni(OH)₂→Ni⁰↓+H₂O+½O₂  (6)

Numerous other steps, (e.g. washing, scrubbing and stripping), arecombined with each of the three main extractions, i.e. Co, Mg and Ni.Altogether about 40 individual mixer-settlers are used, making for quitea complicated and expensive process for Co removal, hence the incentiveto make it simpler.

Finally it is worth emphasizing that the Co extraction by solventextraction in U.S. Pat. No. 6,171,564 is only performed after firstprecipitating a mixed hydroxide, and then re-leaching this solid.

In case it might be wondered why this sequence is required, it is ourcontention that Co extraction by solvent extraction is quite inefficientif such precipitation/releach is not carried out, due to the impuritiespresent in the feed liquor to precipitation.

In other words the Co solvent extraction with Cyanex 272 when applied tothe (acidic) feed liquor to this precipitation is inefficient as itcontains too many interfering impurities, particularly Mg and Mn; bothof these impurities compete with Co in the acidic solutions.

Effectively this means that the hydroxide precipitate has to stay as amixed hydroxide, containing both Ni and Co, as well as Mg and Mn, whichlimits its marketability in practice. It is worth noting that most Nirefineries now use Cyanex 272 for separating Co from Ni, so feedmaterials to these refineries are usually restricted to low Mg and lowMn materials. In practice, this means feed materials to Ni refineriesare currently either Ni mattes or mixed sulphides, not mixed hydroxidesor concentrates; hence the marketability for mixed hydroxide has notbeen established so far, and represents a significant uncertainty forany mine project that depends on selling such a product at a good price.

Reference is also made to patent applications WO 02/22896 A1; WO02/22897 A1; WO 2005/073415 A1; WO 2005/073416 A1 and WO 2006/032097 A1of the Commonwealth Scientific and Industrial Research Organisation(CSIRO) which has been engaged for some years in researching the fieldknown as Synergistic Solvent Extraction (SSX). This technique makes useof two or more extractants combined together to achieve beneficialeffects superior to those of the individual extractants.

Co and/or Ni extraction from solution is the subject of a number ofthese patent applications, but generally they are extracted togetherfrom other impurities. The concept of trying to separate Ni from Co isevidently not contemplated in any of these patent applications.

Also, these patent applications do not contemplate trying to takeadvantage of kinetic differences between Ni and Co in extraction, whichis an important and unusual feature of the present invention, i.e.unusual in solvent extraction.

Rather, extraction efficiencies in these prior applications by CSIRO aregenerally based on steady-state results, i.e. results which approachequilibrium, (the normal situation in solvent extraction).

Patent Application WO 2005/073415 discloses a process for selectivelyextracting Co and/or Mn from leach solutions containing Mg, Ca (andpossibly Mn), using a combination of two organic extractants. Thiscombination is made up of a carboxylic acid such as Versatic™ 10(2-methyl, 2-ethyl heptanoic acid) and a hydroxyoxime such as LIX™ 63(5,8-diethyl-7-hydroxy-6-dodecanone oxime).

The extraction process with this blend has superior selectivitypossibilities for Co, Ni and Mn over the impurities mentioned, comparedto extraction with Versatic 10 alone.

This improvement is expressed as a downward (synergistic) shift in theisotherm for Co, Ni and Mn (the pH₅₀ is reduced by 1-3 pH units). Cu andZn behave similarly (synergistic shift), whereas Mg and Ca behaved inthe opposite sense, i.e. antagonistic shifts to higher pH₅₀.

Taken together, this picture indicates improved separation possibilitiesfor Co, Mn, Ni, Cu and Zn from Mg and Ca. However, it is noted thatextraction of Ni with this organic combination is relatively slowcompared to Co and Mn; Example 2 and FIG. 3 indicate that 10 minutesextraction is needed to get even 85% Ni extraction, (compared to <1minute for Co and Zn), whereas the isotherm (where steady stateconditions apply) indicates almost quantitative Ni extraction ispossible at say pH 5. The implications are that the Ni extractionkinetics aren't fast enough to allow for a practical process for Niextraction, and thus the focus is primarily on Co and/or Mn extraction.

Stripping of Mn and Co from the loaded organic (LO) is accomplishedquickly with dilute acid, presumably at ambient temperatures (notspecified though).

Considering the overall process, (extraction plus scrubbing andstripping), Co can be extracted together with Mn from a leach solution,and thus separated from Mg and Ca; alternatively Co can be extractedpreferentially from Mn as well, presumably by operating at a lower pH.Small concentrations of Mn can also be scrubbed from the LO by Co ifdesired, (presumably this only works if Mn extraction is modest comparedto Co loading).

As explained above, Ni extraction with this system is not attractive, asthe Ni extraction kinetics are too slow to be useful, even though the Niisotherm is similar to that of Co. Some Ni will extract inevitablythough if present in the feed solution, and has to be stripped with Coor separated out subsequent to stripping. So Ni is more of a nuisance toa Co purification process, if anything. If present, Cu and Zn alsoextract in a similar fashion to Co, and must be separated out byselective stripping or in subsequent steps on the strip product stream.

Thus this process is particularly aimed at Co extraction away fromcommon leach liquor impurities, particularly Mg and Ca, where the Co isthe only metal of interest (i.e. Cu, Zn and Ni are missing or in minorconcentrations), and is not particularly attractive for solutions thatcontain high Ni as well as Co.

Patent Application WO 2005/073416 is particularly aimed at Co and/or Niextraction away from common leach liquor impurities, particularly Mn, Mgand Ca. It uses a similar organic mixture as WO 2005/073415, e.g.Versatic 10 and LIX 63, except that a 3^(rd) component is added, aso-called kinetic accelerator like TBP.

This 3^(rd) component literally accelerates the extraction of Ni, sothat both Ni and Co are extracted together, and thus compensates for theperceived slow extraction kinetics of Ni. It also accelerates thestripping of Ni from the organic.

The benefits of the synergistic system together with the accelerator aredescribed in Examples 1-8, and shown in FIGS. 4-9, 11 and 12.

The process has several embodiments, distinguished largely by the designof the stripping circuit, to separate Ni from Co, after the two metalshave been co-extracted into the organic phase. Co strips more easilythan Ni, so selective stripping is an option for separating the twometals, using mild conditions, e.g. with dilute acid or at higher pH.

Thus in Option 1 (page 13), Co is selectively stripped from the loadedorganic, leaving the Ni behind for later stripping under more severeconditions. This option is shown as flowsheets in FIGS. 1-3, anddescribed in Examples 9 and 10.

In Option 2 (page 16), Co and Ni are stripped together, and then thestrip product liquor containing both metals (but notably free from Mg)is then subjected to a 2^(nd) extraction, typically using Cyanex 272,which is selective for Co over Ni, (as noted above) in the absence of Mgand other metals. This option is described in Example 11, and shown inFIG. 10.

Various complications in both options arise with other metals, e.g. Cu,Mn and Zn, which co-extract with Co and Ni, and have to be subsequentlyseparated out during stripping, or by scrubbing, from Co and Ni.

In none of the examples shown, nor in the text or claims, is Coseparated from the Ni in the solution by selective extraction, as in thepresent invention.

SUMMARY OF THE INVENTION

According to the invention, there is provided a process for separatingCo from Ni in an aqueous feed solution, comprising subjecting thesolution containing Ni and Co to extraction with an extractant and usingkinetic differences between Ni and Co in the extraction by controllingduration of the extraction, whereby a major portion of Co and a minorportion of Ni is extracted from the solution to produce a loadedextractant enriched in Co and depleted in Ni compared to the feedsolution and a Co-depleted raffinate containing Ni.

The duration of the extraction may be controlled by limiting theduration of the extraction to a period wherein a maximum ratio of Co:Nihas been extracted or it may be controlled by limiting the duration ofthe extraction to a period wherein a maximum ratio of Ni:Co remainsunextracted or it may be controlled by limiting the duration of theextraction to a period wherein the ratio of Co:Ni that is extracted andthe ratio of Ni:Co that remains unextracted are at optimum values.

The feed solution may be acidic.

The significance of conducting the separation in acid solution, is thatthe process can be applied to a typical leach liquor, (which isfrequently an acid leach of some description). Thus there is no need forprior precipitation of Ni/Co, and subsequent leaching, which has beenrequired until now for efficient Ni/Co separation, as described above.

This invention provides a new and more efficient method for Coseparation, which can be used with fewer prior purification steps, inparticular without the Ni/Co precipitation and releach.

Hence the new process is more rugged than the prior art processes inthis field, i.e. it is able to withstand the presence of key impuritiessuch as Ca, and thus offers a potentially more economical route to Coseparation from Ni.

In addition to Ni, other metals that may be present in the feed solutionare also rejected efficiently into the raffinate, notably Mg and Mn.Thus leach liquors bearing these impurities even in high concentration,can be treated for Co extraction.

The extraction may be carried out at a temperature of about 15° C. to50° C., preferably 20° C. to 35° C. It has been found that a lowerextraction temperature favours the Co:Ni extraction ratio. Even thoughthe Co extraction is less at lower temperatures, the Ni extraction isreduced even further. The optimum temperature may be decided by a numberof factors including ease of operation, number of stages, etc.

The extraction may be carried out at an organic to aqueous ratio ofabout 0.5:1 to 4:1, preferably 0.5:1 to 2:1.

The pH at which the extraction is effected may be controlled by addingan alkali, such as one or more of NaOH, NH₃ and KOH. For example, analkali may be added to the extractant in an amount of about 0.01 to 0.3grammole per litre, preferably 0.02 to 0.2 grammole per litre.

The extraction may be carried out at a pH of about 2.5 to 5.5,preferably about 3.8 to 4.7.

The extraction may be carried out in one or several stages, for exampletwo successive stages. The duration of the extraction may be about 30 to240 seconds per stage, preferably about 30 to 180 seconds per stage.

During a first stage of the extraction, alkali may be added to theextractant at about 0.01 to 0.15 grammole per litre of extractant,preferably about 0.02 to 0.08 grammole per litre of extractant. During asecond stage an alkali may be added to the extractant at about 0.01 to0.15 grammole per litre of extractant, preferably about 0.02 to 0.08grammole per litre of extractant.

The extractant may comprise a mixture of at least two extractants, forexample a carboxylic acid, such as Versatic™ 10 (2-methyl, 2-ethylheptanoic acid), and a hydroxyoxime, such as LIX™ 63(5,8-diethyl-7-hydroxy-6-dodecanone oxime).

The carboxylic acid extractant percentage may be about 2 to 20 v/v %with reference to total volume of the extractants and a diluent,preferably 2.5 to 5 v/v %. The hydroxyoxime extractant percentage may beabout 4 to 40 v/v %, preferably 5 to 30 v/v %.

No kinetic accelerator is used, as in some of the prior art above, andin fact it is an aspect of the invention that the slow kinetics for oneor both Ni extraction and Ni stripping are utilized to facilitate Coseparation, rather than being considered as an impediment to theprocess.

The process may further comprise stripping Co and Ni from the loadedextractant with an acidic strip solution to produce an aqueous Co and Niproduct solution and a stripped extractant which is recycled to theextraction process to complete the cycle. Sufficient acid is used in theacidic strip solution so that the product solution may have a pH ofabout 1.5 to 2.5, preferably 1.6 to 2. The stripping may be carried outat a temperature of about 30° C. to 60° C., preferably about 40° C. to55° C. and may be carried out in one or more stages, for example 1 to 6,preferably 2 to 4, for a duration of about 3 to 15 minutes per stage,preferably 5 to 10 minutes per stage.

The aqueous product solution may be subjected to a further Co extractionstage to produce a Co-loaded extractant, which is subjected to furtherstripping to produce a Co product solution, and a second Ni raffinatewhich is combined with the Co-depleted raffinate from the Ni and Coextraction to produce a Ni product solution.

The further extraction may be carried out with a different extractant,such as Cyanex 272 (bis 2,4,4-trimethylpentyl phosphinic acid).

According to another embodiment, the process may further compriseselectively stripping Co from the loaded extractant from the Ni and Coextraction with a dilute acidic strip solution to produce a Co solutionand a partially stripped extractant and then subjecting the Co solutionto a second Co extraction to produce a Co-loaded extractant and a secondNi raffinate. The Co solution may have a pH of about 1 to 2.5,preferably about 1.7 to 2.2.

The selective stripping may be carried out in 1 or 2 stages for aduration of about 1 to 10 minutes, preferably 3 to 5 minutes, per stageand at a temperature of about 20° C. to 40° C., preferably about 25° C.to 35° C.

The process may further comprise stripping Ni from the partiallystripped extractant with a stronger acidic strip solution to produce aNi solution and a stripped extractant and recycling the strippedextractant to the extraction. The Ni solution may have a pH of about 1to 2, preferably about 1.2 to 1.8.

The Ni-stripping may be carried out at about 30° C. to 60° C.,preferably 40° C. to 55° C. and in one or more stages, for example 1 to6, preferably 2 to 4, for a duration of about 3 to 15 minutes per stage,preferably 5 to 10 minutes per stage.

The process may further comprise scrubbing any Mn present from theloaded extractant, before any stripping, with a scrub aqueous solutioncomprising part of the solution enriched in Co or a part of theraffinate. The Co to Mn ratio in the scrub solution may be about 10:1 to0.75:1, preferably 1.6:1 to 0.4:1. The scrubbing may be carried out at atemperature of about 20° C. to 40° C., preferably 25° C. to 35° C.

Also according to the invention there is provided a hydrometallurgicalprocess for the recovery of Ni and/or Co from an ore or concentrate orother feed material containing Ni and Co, comprising subjecting the feedmaterial to acid leaching to obtain a resultant acid leach solution andsubjecting the leach solution containing Ni and Co to extraction with anextractant and using kinetic differences between Ni and Co in theextraction by controlling duration of the extraction, whereby a majorportion of Co and a minor portion of Ni is extracted from the solutionto produce a loaded extractant and a Co-depleted raffinate containingNi.

Other aspects and features of the present invention will become apparentto those ordinarily skilled in the art upon review of the followingdescription of specific embodiments of the invention in conjunction withthe accompanying Figures and Examples.

BRIEF DESCRIPTION OF THE DRAWINGS

Embodiments of the present invention will now be described, by way ofexample only, with reference to the accompanying drawings, in which:

FIG. 1 is a simplified flowsheet of a process for separating Co from Niin solution shown as being applied in an overall hydrometallurgicalprocess for the treatment of ores or concentrates containing Ni and Co.

FIG. 2 is a flowsheet of one embodiment of the process with bulkstripping of Co and Ni taking place.

FIG. 3 is a flowsheet of another embodiment of the process withselective stripping of Co followed by Ni.

FIG. 4 is a graph showing Co and Ni extraction versus residence time fordifferent organic:aqueous ratios.

FIG. 5 is a graph showing Ni:Co ratio in the raffinate versus residencetime for different organic:aqueous ratios.

FIG. 6 is a graph illustrating the effect of Versatic 10 concentrationon Co extraction.

FIG. 7 is a graph illustrating the effect of Versatic 10 concentrationon Co and Ni extraction at 60 seconds duration.

FIG. 8 is a graph illustrating the effect of LIX 63 concentration on Coextraction.

FIG. 9 is a graph illustrating the effect of LIX 63 concentration on Coand Ni extraction at 60 seconds duration.

FIG. 10 is a graphical illustration showing Mn scrub results asdescribed in Example 7.

FIG. 11 is a graphical illustration showing the stripping kinetics of Coand Ni using 3 and 5 g/l sulphuric acid.

FIG. 12 is a graphical illustration showing the stripping kinetics of Niat varying temperatures.

FIG. 13 is a simplified flow diagram illustrating a bulk strippingprocess with two acid feed streams and two separate product streams asdescribed in Example 16.

FIG. 14 is a graph showing Co:Ni ratio in a loaded extractant forsolvent extractions at different temperatures.

DETAILED DESCRIPTION OF SPECIFIC EMBODIMENTS

With reference to FIG. 1, a process 10 for separating Co from Ni insolution is shown as being incorporated in a hydrometallurgical process12 for the recovery of Ni and Co from an ore or concentrate 13, or otherfeed material containing Ni and Co.

The ore or concentrate 13 containing both Ni and Co is first leached 14to produce a leach slurry 15, from which the leach solution 17 isseparated, as shown at 16, to leave behind a residue 25 by the standardmethods of thickening and/or filtering. Such leach liquors inevitablycontain numerous metals besides Ni and Co, notably Fe, Cu, Zn, Cd, Al,Mg, and Mn, to name a few.

Most of these cause difficulties with the eventual Ni—Co separation,(using known technology, as discussed above), so must be dealt withfirst. This is effected in a purification stage 18 where some of theseimpurities 26, such as Cu, Zn, Fe and Al, are separated from the Ni—Cosolution 17 obtained from the leaching stage 14 to result in the aqueousfeed solution 19. The purification stage 18 invariably consists of anumber of steps, including solvent extraction, precipitation ofimpurities, etc. The details of the purification process 18 naturallydepend on the impurities present and their relative concentrations.

Thus if Cu is present in economically significant concentrations,solvent extraction would be considered as this is an effective andselective process for Cu recovery. If only trace amounts of Cu arepresent, sulphide precipitation would likely be used, as this alsoremoves Cd and Zn very effectively in a Ni/Co solution. Fe and Alremoval generally is accomplished by neutralization with limestone orlime. Mg and Mn removal is more problematic as has been discussedpreviously. However, in the present process removal of Mg and Mn, alongwith Ca, is not required as these impurities are effectively controlledduring the extraction.

The next step in the process is Co extraction 20, the duration of whichis controlled so that the bulk of the Co in the solution 19 is extractedalong with a minor portion of the Ni, as will be described in greaterdetail below, to obtain a Co—Ni solution 27 from which Co metal productcan be recovered. The raffinate 21 from the solvent extraction 20 is apurified Ni solution.

The following step is Ni recovery 22 from the raffinate 21 to finalproduct 24, e.g. by electrowinning. The raffinate 28 from the Nirecovery 22 is recycled to the leaching stage 14.

The invention has several embodiments, largely dependent on howstripping is effected after Co extraction 20. Essentially Co can beselectively stripped with respect to Ni, and then the Ni strippedsubsequently, or Co and Ni can be stripped together in a bulk strippingoperation.

In addition there are other embodiments of the invention suited forvarying feed materials, such as solutions derived from leaching lateriteore, instead of sulphide concentrates. Such solutions may have differentNi:Co ratios and different levels of impurities (Mg and Mn for example),and thus require a modified flowsheet for optimum results.

In the process, there are two basic steps in the solvent extractionprocess: extraction and stripping. In addition, a third step,saponification, is sometimes needed as a separate step beforeextraction, or is sometimes combined with extraction.

The extraction process 20 typically uses a mixture of two or moreextractants and critically the time of extraction is controlled so as tooptimize Co extraction at the expense of Ni extraction.

Co extraction is typically quite fast compared to Ni extraction, withretention times measured in seconds rather than minutes. Ni extractionis never zero, but by careful control of retention times, as well asother factors, the Co extraction can reach almost quantitative levels,with less than 20% Ni extraction. This enables the process toeffectively purify the Ni solution of its Co content, and reach a Ni:Coratio of >667:1, sufficient to satisfy the Co specification for LMEGrade Nickel. (Co specification is 0.015% maximum).

In the saponification step, the organic stream is contacted with analkali such as caustic to pre-neutralize, or saponify, the organicextractant prior to contact with the aqueous stream. Alkali addition isrequired during or before extraction because the extraction processcreates acid, and if not compensated, the acid will lower the pH of thefeed solution. Co extraction is dependent on the pH, typically it shouldbe maintained in the pH 2.5-pH 5 region (this will be discussed in moredetail below), so to avoid acidification and a consequent drop in pH,alkali should be added, either before or during extraction or both.

In extraction, cobalt and some nickel are extracted from a low acidaqueous solution. The extracted cobalt and nickel are transferred(loaded) onto the organic stream with a reaction analogous to reaction(3) above. Reference to “organic stream” or “organic” implies theextractant with a diluent. Therefore, where more than one extractant isused, the term implies a mixture of the extractants and a diluent.

The extraction process 20 typically takes place in more than one stage,wherein the aqueous feed solution is contacted by organic streams morethan once. There are several possible arrangements for such amulti-staged process and counter-arrangements are sometimes used,wherein the organic feed stream and the aqueous feed stream move indifferent directions.

A preferred embodiment is a series parallel extraction circuit, whereinthe aqueous feed solution is contacted by successive increments oforganic stream, each one being a freshly stripped organic. A typicalarrangement for the extraction is shown in FIG. 2, with a seriesparallel circuit. FIG. 2 will be described in more detail below. Thefeed material to the process shown in FIG. 2 is the aqueous feedsolution 19, referred to in FIG. 1, containing Ni and Co, a solutionthat has already been treated to remove certain impurities, such as Cu,Zn, Cd, Fe, Al, for example, as described above.

Lower temperatures of extraction favour the Co:Ni extraction ratio, dueto the more rapid decrease in Ni extraction at low temperature, as seenin Table 38 and FIG. 14. FIG. 14 shows the effect of temperature on theCo:Ni ratio obtained in extraction with the new invention; tests wererun at pH 4.5, with an organic to aqueous ratio of 1:1, using 4.7 v/v %Versatic 10 and 14 v/v % LIX 63 in Shellsol D70 as the extractant. Theextraction was from an aqueous feed solution with a Ni:Co ratio ofapproximately 20:1.

As can be seen from Table 38 at 22° C. the Co extraction is slightlylower than at higher temperatures, but the Ni extraction is much lowerat 22° C. compared to say 40° C., leading to a greatly increased Co:Niratio in the loaded extractant.

Prior to the extraction, saponification is effected, wherein thestripped organic is neutralized with an alkali, as discussed above.

In another embodiment, the saponification takes place simultaneouslywith the extraction, at least in part, so some alkali is added to theorganic before extraction, and some is added into the mixture of aqueousand organic during extraction, to maintain a target pH, as will bediscussed in the Examples below.

The final step is stripping, where the cobalt and nickel are strippedoff the organic stream with acid and released into an aqueous productsolution. The bulk stripping option is shown in FIG. 2, and the strippedorganic stream is then recycled back to saponification and/orextraction, to complete the loop.

In FIG. 2 with bulk stripping, the strip liquor typically will containan approximate nickel to cobalt ratio of about 3:1, depending on theNi:Co ratio in the feed solution 19.

A more detailed description of FIG. 2 will now be given showing aflowsheet of one embodiment 30 of the invention where bulk stripping ofCo and Ni takes place.

In this embodiment, the extraction 20 takes place in two stages, namelya first extraction 31 and a second extraction 32. The organic stream isshown in broken lines and the aqueous stream is in solid lines.

The extractions 31 and 32 use a combination of two extractants toextract most of the Co in aqueous feed solution 19.

The extraction 31 produces an intermediate raffinate 50 which issubjected to the second extraction 32 which in turn produces a raffinate51 containing the unextracted Ni and impurities such as Ca, Mg and orMn.

The loaded organic (LO) 49 and 42 from the extractions 31 and 32,respectively, are combined as shown at 77 to produce a combined loadedorganic 34.

For pH control, alkali 33 is added to extractions 31 and 32, either bybeing added to the organic prior to contact with the aqueous feed 19, 50(saponification) or during the extractions 31 and 32, or both.

To summarize, the products of the extractions 31 and 32 are an aqueousstream (raffinate) 51 denuded of Co as much as possible and a loadedorganic (LO) 34 containing the extracted Co values. As discussed above,this is achieved by careful control of the duration of the extractions31 and 32.

Approximately 97% to 98% of the Co is extracted with a minimum of Nico-extracted, typically 10 to 20%.

The loaded organic 34 is stripped of the Co and Ni by subjecting it tostripping which is carried out in three stages 110, 111 and 112, incounter-current mode.

This is effected by feeding an acidic stream 37 to the third stage 112of the stripping operation and feeding the loaded organic 34 to thefirst stage strip 110. The organic 34 and the acidic stream 37 flow inopposite directions through the stripping stages 110, 111 and 112,wherein Co and Ni are stripped from the organic 34 during each of thethree stripping stages 110, 111 and 112. This produces a strippedorganic (SO) 40 which is recycled to the extractions 31 and 32 and astrip product 38 which contains virtually all of the Co values, togetherwith some Ni.

The conditions of stripping in 110, 111 and 112 are controlled tomaximize Co stripping but not the Ni stripping. For economic reasons itis useful to recover nearly all of the Co in the loaded extractant 34 bystripping. However, it is neither necessary nor advantageous tocompletely strip Ni.

Although, it is beneficial to strip most of the Ni from the loadedextractant 34 to allow the next cycle of extraction to proceedefficiently, a small amount of Ni left behind on the stripped extractant40 allows the stripping to be carried out in slightly less severeconditions (less acid used for example), and this is still compatiblewith efficient stripping in the next extraction cycle. Starting with say1 to 2 g/l Ni in loaded organic, it is found that about 200-250 mg/l[Ni] may be left on the stripped organic stream 40, and still achievesatisfactory extraction with the next cycle.

This allows the stripping to be carried out with the minimum of acidconsumed, and also produces a strip product 38 with minimum acidity,which benefits future processing options.

Typically the Ni:Co ratio in the strip product 38 is about 3:1, due tothe predominance of Ni in the original feed solution 19, but this variesof course according to the ratio in the feed solution 19.

The strip product 38 is now further refined by a further Co extraction36, which completes the Co separation process, and returns the smallamount of Ni back to the main Ni-bearing stream, for eventual Nirecovery. It is now possible to use known processes for the extraction36, because the first extractions 31 and 32 have effectively eliminatedmost of the damaging impurities, such as Mn, Mg and Ca, which wouldotherwise make this uneconomic. This point is further illustrated in theExamples below. The extraction 36 with Cyanex 272 as extractant is nowvery selective for Co compared to Ni, and virtually no Ni is extracted,thus producing a very high Co:Ni ratio in the organic stream (LO) 42. Inthis case alkali 33 is also added for pH control.

Any traces of Cu or Zn in the original feed solution 19 will report tothe strip product 38 and can now be separated from the Co in the Coextraction 36 as is commonly done in Co refineries.

The loaded organic 42 from the extraction 36 is subjected to stripping44 with acid 39 to produce a strip product (Co solution) 46 and strippedorganic 48 which is recycled to the extraction 36.

The strip product 46, which is essentially a pure Co stream, free of Niand other impurities, can be easily processed to a marketable Coproduct, such as solid CoCO₃, by precipitation with NaCO₃, a method wellknown in the industry.

The raffinate 51 from the second extraction 32, which has a very highNi:Co ratio, is combined with the raffinate 52 from the extraction 36,as shown at 54, to produce a combined raffinate 55. The raffinate 55 istreated for Ni-recovery by various methods well known in the industry.

The Ni raffinate 55 produced is suitable for a number of subsequentsteps; for example Ni may be precipitated by MgO as a hydroxide with anickel to cobalt ratio of at least 667:1.

Another embodiment 60 of the process in which selective stripping of Cotakes place is shown in FIG. 3. Steps that correspond with theembodiment 30 of FIG. 2 are given similar reference numerals. Again, theorganic stream is shown in broken lines.

The Co and Ni from the aqueous feed solution 19 is subjected to the twosolvent extractions 31 and 32 arranged in a series parallel extractioncircuit. Again in this case, alkali 33 is added for pH control.

In the first extraction 31 the solution 19 is contacted with strippedorganic 47 producing a loaded organic 49 and an aqueous stream(intermediate raffinate) 50 with most of the Co extracted from it.

The intermediate raffinate 50 is then subjected to the second extraction32 where it is contacted with further stripped organic 47 to extract theremaining Co and produce a raffinate 52 with the desired Ni:Co ratioabove 667:1. Again the duration of the extractions 31 and 32 iscontrolled, as described above.

The loaded organic 49 from the first extraction 31 is combined with theloaded organic 42 from the second extraction 32 and subjected to a Costrip 62 with dilute acid 64 designed to selectively strip Co from theloaded organic at the expense of Ni by means of mild conditions, i.e.low acid, low temperature and short retention time. Ni requires moresevere conditions to be efficiently stripped, and thus it is possible toselectively strip most of the Co with minimum Ni, by choosing mildconditions.

This is achieved by careful control of the amount of acid in the diluteacidic strip solution, so that the stripping conditions have the minimumnecessary acidity for efficient Co stripping. In this way, Co isselectively stripped with respect to Ni.

The Co product solution 66 coming from 62 may have a pH of about 1 to2.5 preferably about 1.7 to 2.2.

The partly stripped organic 68 now depleted in Co, is then subjected toa Ni strip 74 with strong acid 76 under conditions suitable forNi-stripping, producing the stripped organic 47 in which the [Ni] isreduced to about 200 to 250 mg/l. The stripped organic 47 is recycled tothe extractions 31 and 32, completing the loop.

As before in the bulk stripping mode, it is not necessary to strip allof the Ni from the organic stream 68, and a certain residual level of Niis deemed to be economically beneficial to the process.

The conditions of the Ni stripping are controlled to strip most but notall of Ni from the loaded extractant, in the same fashion as describedfor the previous embodiment 30, shown in FIG. 2, i.e. to achieve atarget of about 200-250 mg/l [Ni] in the stripped extractant.

The aqueous product (raffinate) 66 from the Co strip 62 is purified ofits Ni content by a further extraction 70, using Cyanex 272 or someother suitable Co extractant, producing a Ni containing raffinate 78.Alkali 33 is again added for pH control.

The loaded organic 80 from the extraction 70 is stripped, as shown at72, with dilute acid 64 as before to produce a pure Co solution 82.

The stripped organic (SO) 120 from the strip 72 is split, as shown at122, into a first stream 124 which recycled the extraction 70 and asecond steam 126.

The aqueous product (raffinate) 84 from the Ni strip 74 is subjected toCo extraction 86 with the second organic stream 126 to remove residualCo, producing a loaded organic 88, which is recycled to the strip 72 forextracting Co therefrom. The solution 82 is a Co product solution 96which is very low in Ni and other impurities and therefore suitable forCo recovery, such as with CoCo₃ precipitation or some other knownprocess.

The aqueous solution (raffinate) 100 from the Co extraction 86 iscombined as shown at 102 with the raffinates 52 and 78 to produce acombined Ni raffinate 104 which is low in Co and therefore suitable forsubsequent Ni recovery. These methods include as a first stepprecipitation as either a hydroxide (Ni hydroxide) or sulphide (NiS),which are then readily refined.

A number of known processes, including pyrometallurgical (smelting) andhydrometallurgical techniques (both acidic and alkaline) have been usedon a commercial scale for this purpose. With Co already removed theremaining refining process for Ni is greatly simplified.

The invention is based on the discovery that Co can be extractedselectively from Ni, using a suitable extractant, such as a mixture ofextractants, namely a carboxylic acid and a hydroxyoxime, and by takingadvantage of faster extraction kinetics for Co compared to Ni.

Specifically, by using short retention times, Co can be very efficientlyextracted, to achieve >97% Co extraction, whilst minimizing Niextraction, i.e. <20% extraction.

Thus leach liquor resulting from leaching a typical Ni—Co sulphideconcentrate with a Ni:Co ratio of 20:1, can be refined by the process ofthe invention to produce a Ni raffinate with Ni:Co ratio of about 667:1.This ratio is sufficient to produce a Ni product that satisfies LMEGrade requirements for refined nickel metal, with respect to Co content.

The invention will now be further described with reference to thefollowing examples. A summary of the Examples is given in Table 1.

TABLE 1 List of Examples Example # Purpose Conclusion 1 Limitations ofknown technology High amounts of Mg and Mn co-extraction, leading forextracting Co from a sulphide to extremely high reagent consumptionleach liquor containing Mg, Mn 2 As in example 1, but with more Evenlarger amounts of Mg and Mn co-extraction, impurities present, (lateriteleach leading to even worse reagent consumption liquor) 3 Batch modeexample of the new 76% Co extraction could be achieved in one stageinvention on leach liquor, in one with only 5% Ni extraction, andnegligible Mg or Mn stage extraction, with low reagent consumption 4Batch mode example of Two stages extracted >98% of Co with about 12%invention using two stages of Ni leading to a final Ni:Co in raffinateat over 1000, extraction well above the target. 5 Optimizing the organicLower Versatic 10 and higher LIX 63 concentrations composition, i.e.ratio of the two produce higher Co extractions with a slight extractantsreduction in selectivity 6 Treatment of laterite leach liquor Produces aNi:Co ratio of 805 with low reagent with new invention consumption withonly minor Mn extraction into the loaded organic stream 7 Mn Scrub forlaterite leach Mn can be successfully scrubbed from the loaded liquors.organic by a dilute Co stream at high efficiency 8 Comparison of testingmethods, Continuous operations continually produced equipment selectionand raffinate solutions >667:1 Ni:Co with saponified neutralizationmethods organic and a pre-mix stage combined with a pipe reactor 9Continuous Operations for Continuous operations over a one month periodSulphide Feed showed the new invention was capable of producing LMEGrade solution continuously. 10 Continuous Operations for Continuousoperations continuously produced LME Laterite Feed Grade raffinate 11Selectively strip Co from loaded The majority of the Co can beselectively stripped at organic low acid concentrations, short retentiontimes and low temperature 12 Stripping Ni from the partially Ni can bestripped to low levels at elevated loaded organic temperatures (50° C.)13 Continuous Co selective The majority of the Co can be selectivelystripped at stripping utilizing a dilute strip low acid concentrationand O:A ratios of 5:1 solution or a concentrated strip solution 14Continuous Ni stripping of Ni can be stripped to <0.25 g/L Ni inorganic. partially loaded organic 15 Bulk stripping of loaded organicBulk stripping can be utilized to produce a final strip product solutionwith <3:1 Ni:Co 16 Continuous operations of new Co extraction wassuccessfully achieved, two acid invention with two acid feed feeds werenot successful as too much Co was stripping present in the stripsolution to combine with final raffinate

Example 1

This example illustrates the limitations of using existing technologyfor extracting Cobalt from Nickel-Cobalt sulphide leach liquor, due tothe effect of impurities.

A synthetic leach liquor (PLS) as shown in Table 2 was prepared fromsulphate salts, and was intended to simulate leach liquor derived fromleaching a typical Ni sulphide concentrate. It was extracted in a singlestage with bis 2,4,4-trimethylpentyl phosphinic acid, (i.e. thecommercial reagent Cyanex 272); the organic phase was made up to 10 v/v% Cyanex 272 with Shellsol 2046AR as diluent.

The objective was to extract sufficient Co from the leach liquor so thatthe resultant raffinate met the Co requirements for LME Grade Nickel,(Ni:Co ratio >667:1). In effect this meant >98% Co extraction wasrequired, depending on the extent of Ni extraction. Because someextractant may be consumed by impurities (i.e. not by Co alone), theamount of extractant needed is not easily predicted; thus to cover allreasonable eventualities, four different extraction tests were carriedout using different O:A ratios as shown in Table 3.

Nickel Cobalt and Impurity Extraction

The organic phase was placed in a one litre stainless steel rectangularbox immersed in a water bath at 60° C. An overhead stirrer was used formixing, and the required volume of aqueous solution was slowly added tomaintain an organic continuous continuity and left to mix for 5 minutes.The pH was measured regularly and adjusted by addition of 50% sodiumhydroxide to maintain pH 5.1-5.5, (the pH range that is understood to beused in commercial plants for Cyanex 272).

After 5 minutes of mixing under these conditions, the resulting emulsionwas separated in a one litre separatory funnel, producing raffinate andloaded organic. Feed and product assays for each ratio are presented inTable 2.

TABLE 2 Feed and Product Assays Example 1 Aqueous Organic O:A Ni Ca ClCo Mg Mn Na Ni Ca Co Mg Mn Na Ratio ppm ppm ppm ppm ppm ppm ppm ppm ppmppm ppm ppm ppm PLS/SO 19510 468 7130 994 9984 474 1522 0 0 0 0 0 0 5:118710 454 7060 14 2751 2 16170 46 11 191 1419 100 3 3:1 19240 485 722047 5812 9 11190 23 7 295 1295 152 1 1:1 19650 525 7010 165 8300 32 628927 3 807 1749 460 0.5 0.5:1   19770 524 7030 335 9213 69 4484 40 3 12701736 846 0.5

Percent extraction of each element for each ratio is presented in Table3.

TABLE 3 Extraction Results Example 1 O:A Ni Ca Co Mg Mn g NaOH:g RatioNi:Co % % % % % Co 5:1 1336 0.2 3.0 98.6 72.4 99.6 25.4 3:1 409 0.1 0.095.3 41.8 98.1 17.9 1:1 119 0.1 0.0 83.4 16.9 93.2 8.97 0.5:1   59 0.20.0 66.3 7.7 85.4 8.78

Caustic consumption in each test was calculated from the aqueous Na andCo assays, and is shown as a ratio to the grams Cobalt extracted. Allextractions are based on the aqueous solutions with the exception of Niwhich is based on organic assay.

Conclusions

Co was extracted in all cases without any Ni extraction, butunfortunately both Mg and Mn co-extracted with the Co. Using a largeexcess of organic (5:1 ratio), the raffinate did finally meet therequired target for Ni:Co ratio. However, much of the Mn and Mg werealso loaded at this point. Consequently reagent consumption during this(5:1) test was very high at 25.4 g of NaOH/g Co which probably makes useof this reagent uneconomic for such leach solutions.

This example illustrates the fundamental difficulty facing theseparation of Co from Ni in the presence of common impurities such asMg, Mn and Ca, found in leach liquors.

Example 2

This example again illustrates the difficulties of using existingtechnology for extracting Cobalt from Nickel, this time from lateriteleach liquors, which tend to have higher impurities than sulphide leachliquors.

Nickel, Cobalt and Impurity Extraction

A synthetic PLS as shown in Table 4 was prepared as before from sulphatesalts, and was intended to simulate leach liquor derived from leaching atypical Ni laterite. Even with a 5:1 O:A ratio using 10 v/v % Cyanex 272as before, the results were still short of the target (too much Co leftin raffinate), so it was subjected to another stage.

This test was completed as a two stage series parallel test. The firststage consisted of contacting fresh organic solution and PLS solution ina similar arrangement as in Example 1. After 5 minutes under theseconditions, the resulting emulsion was separated, producing raffinateand loaded organic streams.

The raffinate was still not at the target Ni:Co ratio (667:1) so wasthen contacted again in a 2^(nd) Stage with fresh organic under the sameconditions, but with an O:A ratio of only 1:1.

Feed and product assays for both stages are presented in Table 4.

TABLE 4 Feed and Product Assays Example 2 Aqueous Organic O:A Ni Ca ClCo Mg Mn Na Ni Ca Co Mg Mn Na Stage Ratio ppm ppm ppm ppm ppm ppm ppmppm ppm ppm ppm ppm ppm PLS/SO 7671 196 8610 726 12667 5,216 565 0 0 7260 0 0 1 5:1 7086 185 8410 69 8905 162 11451 7 2 69 584 826 2 2 1:1 7291182 8480 10 6750 7 15460 16 1 10 98 148 2

Percent extraction of each element for each stage is presented in Table5.

TABLE 5 Extraction Results Example 2 O:A Ni Ca Co Mg Mn g NaOH:g StageRatio Ni:Co % % % % % Co 1 5:1 103 0.1 5.6 90.5 29.7 96.9 36.2 2 1:1 7290.2 7.1 98.6 46.7 99.9

All extractions are based on the aqueous solutions with the exception ofNi which is based on organic assay.

Conclusions

Co was again extracted preferentially over Ni, and with two stages and alarge excess of organic the raffinate did finally meet the requiredtarget for Ni:Co ratio. However, virtually all of the Mn and asignificant quantity of Mg were also loaded at this point. Reagentconsumption during this test was calculated (from Na assays) at 36.2 gof NaOH/g Co which is very high, and indicates the difficulty with useof this reagent for such laterite leach liquors.

This example again illustrates the fundamental difficulty facing theseparation of Co from Ni in the presence of common impurities such asMg, Mn and Ca, found in leach liquors.

Example 3

This example illustrates the use of the invention for extracting cobaltfrom nickel-cobalt sulphide leach liquor complete with impurities suchas Mg and Mn in substantial concentrations.

A synthetic leach liquor (PLS) as shown in Table 6 was prepared fromsulphate salts. It was extracted in one stage with an organic phasecontaining 4.7 v/v % Versatic 10 and 14 v/v % LIX 63 with Shellsol D80as diluent. (In this description, v/v % is with reference to the totalvolume of the extractants and the diluent.) As in Examples 1 and 2, theobjective was to extract sufficient Co to achieve a Ni:Co ratio >667:1in raffinate. Two separate tests were done, with different O:A ratios,1:1 and 0.5:1.

Nickel, Cobalt and Impurity Extraction

The organic and aqueous phases were mixed in a one litre stainless steelrectangular box immersed in a water bath at 30° C. for one minute at pH4.5. Samples were taken at 15 second intervals, and analyzed for Ni andCo in both phases.

TABLE 6 Feed Assays Example 3 Aqueous Organic Ni Ca Cl Co Mg Mn Na Ni CaCo Mg Mn Na Stream Ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppmPLS/SO 20288 561 11485 1024 11527 425 256 0 0 0 0 0 0

The results are presented in Table 7 and Table 8 and graphically in FIG.4 and FIG. 5.

TABLE 7 Comparison of the Effect of O:A Ratios on Co and Ni ExtractionExample 3 Aqueous Organic Time Ni Co Na Ni Ca Co Mg Mn Na O:A Ratio sppm ppm ppm ppm ppm ppm ppm ppm ppm 1 15 19558 768 595 229 4 249 5 2 330 19309 599 840 366 2 428 3 1 3 45 19165 448 1093 535 2 583 4 2 4 6018871 339 1299 668 1 695 3 1 5 90 18607 242 1648 988 2 798 6 <1 4 18018033 463 1844 1487 3 553 3 <1 3 0.5 15 20250 861 484 315 1 347 1 3 5 3019704 718 679 562 1 705 2 2 4 45 19377 572 942 785 1 876 5 3 4 60 19285557 1070 1071 2 895 5 1 4 90 19137 660 1067 1242 1 687 4 <1 3 180 18744921 1115 1862 1 163 2 <1 2

TABLE 8 Percent Extraction Example 3 Ni:Co O:A Time Raffi- Co:Ni Ni CaCo Mg Mn g NaOH: Ratio s nate Organic % % % % % g Co 1 15 25 1.09 1.10.71 25.0 0.04 0.47 2.3 30 32 1.17 1.8 0.36 41.5 0.03 0.24 2.4 45 431.09 2.6 0.36 56.3 0.03 0.47 2.5 60 56 1.04 3.3 0.18 66.9 0.03 0.24 2.690 77 0.81 4.9 0.36 76.4 0.05 0.0 3.1 180 39 0.7 7.3 0.53 54.8 0.03 0.04.9 0.5 15 24 1.10 1.6 0.18 15.9 0.01 0.71 2.4 30 27 1.25 2.8 0.18 29.90.02 0.47 2.4 45 34 1.12 3.9 0.18 44.1 0.04 0.71 2.6 60 35 0.84 5.3 0.3645.6 0.04 0.24 3.0 90 29 0.55 6.1 0.18 35.5 0.03 0.0 3.9 180 20 0.09 9.20.18 10.1 0.02 0.0 14.5

All extractions were based on the organic assays with the exception ofCo which was based on aqueous assays.

The Co extraction was very fast and peaked at about 50 seconds at anO:A=0.5, or at 90 seconds with O:A=1.0. Ni extraction was significant,but was low in comparison to Co, so that the Ni:Co ratio in theraffinate improved by about a factor of 3.5 at the optimum time.However, this was quite insufficient to achieve the target in raffinate.

At an O:A ratio of 1:1, the maximum Ni:Co ratio in the raffinateoccurred at 90 seconds while the maximum Co:Ni ratio in the organicsolution occurred at about 30 seconds. At an O:A ratio of 0.5:1, themaximum Ni:Co concentration in the raffinate occurred at about 60seconds while the maximum Co:Ni ratio in the organic solution occurredat about 30 seconds. This means that it is necessary to find the bestcompromise between the Ni:Co ratio in the raffinate and the Co:Ni ratioin the loaded organic.

Conclusion

The extraction of Co with the new process is very selective compared toNi.

Co loaded quickly onto the organic until a given time whereas the Nicontinued to load (with additional time) and eventually scrubbed the Cooff the organic. However, the best Ni:Co ratio in the raffinate wasstill <80:1, a long way from the target of 667:1, but this was just withone stage of extraction. This example therefore illustrates thepreference for multiple stages of extraction to achieve the desiredNi:Co ratio in the raffinate, which is achieved in the next example.

Example 4

This example illustrates another embodiment of the new process forextracting Cobalt from Nickel-Cobalt sulphide leach liquor, specificallythe use of multiple stages of extraction to achieve a greater % Coextraction, and thus increase Ni:Co ratio in raffinate.

A synthetic leach liquor (PLS) as shown in Table 9 was prepared fromsulphate salts, and was intended to simulate leach liquor derived fromleaching a typical Ni sulphide concentrate. It was extracted in a twostage series parallel configuration with an organic phase that was madeup to 4.7 v/v % Versatic 10 and 14 v/v % LIX 63 with Shellsol D80 asdiluent. In series parallel mode the PLS is contacted with stripped orfresh organic resulting in a raffinate and loaded organic. The raffinateproduced from the first stage of extraction is then contacted withstripped or fresh organic resulting in a final raffinate and a loadedorganic. The individual product organics from both stages are combinedfor further processing i.e. stripping.

As in Example 3, the objective was to try to extract sufficient Co fromthe leach liquor to meet Co requirements for LME Grade Nickel. In thisexample two stages were used, and given the importance of the relativekinetics of Co extraction vs Ni, the effect of the retention time in thesecond stage was looked at in detail.

Nickel, Cobalt and Impurity Extraction

The organic and aqueous phases were mixed in a 1:1 O:A ratio for 60seconds, using the same procedure as in Example 3, for the first stageof extraction. The resulting Co-depleted aqueous solution was thencontacted with fresh organic solution again for another 90 seconds at anO:A ratio of 1:1 and sampled at various intervals. The pH was measuredand regularly adjusted using sodium hydroxide. Product assays andpercent extraction numbers are presented in Table 10 and Table 11.

TABLE 9 Feed Assays Example 4 Aqueous Organic Ni Ca Cl Co Mg Mn Na Ni CaCo Mg Mn Na Stream ppm ppm Ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppmPLS/SO 19573 551 11700 1011 12649 419 270 0 0 0 0 0 0

TABLE 10 Product Assays Example 4 Aqueous Organic Time Ni Co Na Ni Ca CoMg Mn Na Stage s ppm ppm ppm ppm ppm ppm ppm ppm ppm 1 60 18134 252 1485789 1 754 1 7 3 2 15 17611 124 1861 358 <1 132 <1 7 <1 30 17596 71 2157621 <1 192 <1 8 <1 45 16941 44 2261 826 <1 216 <1 3 <1 60 16831 26 25231067 <1 232 <1 <1 4 90 16477 15 2862 1404 <1 249 <1 <1 19

TABLE 11 Percent Extraction Example 4 Time Ni: Ni Ca Co Mg Mn g NaOH:Stage s Co % % % % % g Co 1 60 72 4.2 0.2 75.0 0.0 1.7 2.8 2 15 142 2.00.0 51.5 0.0 1.7 3.1 30 248 3.4 0.0 73.1 0.0 1.9 3.5 45 385 4.6 0.0 83.20.0 0.7 3.6 60 647 6.0 0.0 90.1 0.0 0.0 4.0 90 1099 7.8 0.0 94.5 0.0 0.04.5

All extractions are based on the organic solutions with the exception ofCo which is based on aqueous assay.

Conclusions

98.5% Co and 12.0% Ni were extracted in two stages with minimalco-extraction of unwanted impurities; as a result the reagentconsumption was only 4.0 g of NaOH being consumed per g of Co extracted.This consumption of NaOH is >6× less than that observed in Example 1using existing technology (for a similar solution), to achieve thetarget Ni:Co in the final raffinate.

This example illustrates how this invention efficiently separates Cofrom Ni in Ni sulphide leach liquors, and thus achieves a Ni:Co ratiosufficient for LME Grade Nickel.

Example 5

This example illustrates the use of the new process for extractingCobalt from a Ni sulphide leach liquor, specifically the use of variedorganic compositions to achieve a greater % Co extraction, and Co:Niratio in the organic.

The organic phase for each test was made up with LIX 63 and Versatic 10with Shellsol D70 diluent.

The objective was to try to determine the effects on Co and Niextraction kinetics by varying the concentration of only one extractantat a time.

Nickel and Cobalt Extraction

TABLE 12 Effect of Organic Composition for Example 5 Feed Organic LoadedOrganic Extraction Versatic LIX 63 Co Ni Co Ni Co:Ni 10 v/v % v/v % g/Lg/L % % Ratio 4.73 6.98 0.40 0.61 20.0 1.52 0.66 4.73 13.96 0.69 0.6767.2 3.42 1.03 4.73 20.94 0.81 0.81 81.4 4.06 1.00 1.70 13.96 0.71 0.5968.5 2.90 1.20 2.37 13.96 0.67 0.56 65.3 2.80 1.20 3.41 13.96 0.74 0.6472.0 3.20 1.15 4.73 13.96 0.74 0.71 72.3 3.50 1.04 7.10 13.96 0.71 0.8068.9 4.00 0.89

Co extraction varied slightly with changing Versatic 10 concentration inthe range of 1.70-7.10 v/v %.

An increasing trend can be observed for Ni extraction with increasingVersatic 10 concentrations resulting in a decreasing trend in the Co:Niratio in the organic solution which suggests that a lower Versatic 10concentration favours a higher Co:Ni ratio. See FIGS. 6-9.

Both Co and Ni extractions increased with increasing LIX 63concentrations from 6.98-20.94 v/v %. The Co extraction increased from20% to 81% with a subsequent increase in Ni extraction from 1.5% to4.06%.

FIG. 9 shows that increased LIX 63 concentration has the same effect onCo and Ni extraction with the selectivity of the Co:Ni in the organicslightly decreasing from 13.96-20.94 v/v %.

Conclusions

A Versatic 10 concentration of 3.41% yielded the highest Co:Ni ratio andthe highest Co extraction at 60 s. An increase in LIX 63 concentrationalso increased the Co and Ni extraction with a slight decrease in theselectivity.

Example 6

This example illustrates the use of the new process for extractingCobalt from a Ni—Co laterite leach liquor.

A synthetic leach liquor (PLS) as shown in Table 13 was prepared fromsulphate salts, to simulate a leach liquor derived from leaching atypical Ni laterite, with increased Mg and Mn in solution, as well as ahigher Co:Ni ratio, i.e. about 1:10 instead of 1:20 with sulphide leachliquor.

As in Example 4, it was extracted in a two stage series parallelconfiguration with an organic phase that was made up to 3.8 v/v %Versatic 10 and 19.5 v/v % LIX 63 with Shellsol D80 as diluent.

Nickel, Cobalt and Impurity Extraction

The organic and aqueous phases were mixed in a 1:1 O:A ratio, using thesame procedure as in Example 4, with retention times of 120 seconds forthe first stage of extraction and 151 seconds for the second stage ofextraction. The pH was measured and regularly adjusted using sodiumhydroxide, as before.

TABLE 13 Feed Assays Example 6 Aqueous Organic Ni Ca Cl Co Mg Mn Na NiCa Co Mg Mn Na Stream ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppmppm PLS/SO 6980 212 2632 696 12200 4980 491 251 2 4 <1 <1 <1

Product assays and extraction values for each stage are presented inTable 14 and Table 15.

TABLE 14 Product Assays Example 6 Aqueous Organic Time Ni Co Na Ni Ca CoMg Mn Na Stage s ppm ppm ppm ppm ppm ppm ppm ppm ppm 1 120 6922 120 1199604 1 602 1 78 64 2 151 6442 8 1878 763 2 125 2 238 78 Combined LO — — —— 685 2 365 1 158 74

TABLE 15 Percent Extraction Example 6 g NaOH/ Ni Ca Co Mg Mn g Co Stage% % % % % extracted E1 5.1 0.0 82.8 0.0 1.6 3.3 E2 7.4 0.0 93.3 0.0 4.8Overall 12.5 0.0 98.9 0.0 4.8

All extractions are based on the organic assays with the exception of Cowhich is based on aqueous assay.

Conclusions

Co was preferentially extracted compared to Ni from the laterite leachliquor; despite the higher ratio of Co:Ni and the increased impuritylevels in this example, the resulting raffinate again exceeded thetarget Ni:Co ratio, achieving 805:1.

A minor amount of Mn was co-extracted with the Co which is potentialproblem for a Co product. However in another embodiment of thisinvention it is shown that Mn on the loaded organic can be separatedwith a Mn scrub stage, before stripping.

The caustic consumption during this test was 3.3 g NaOH/g Co extracted,which is very low (10%) compared to the comparable test using existingtechnology, (Example 2). This example again illustrates how theinvention allows for a more efficient separation of Co from Ni in spiteof an increased level of impurities as found in laterite leach liquors.

Example 7

This example illustrates another embodiment of the invention,specifically the use of a Mn scrub stage to remove Mn from loadedorganic resulting from extraction of laterite leach liquors.

The combined loaded organic used for this test was taken from theextraction circuits in a continuous pilot plant. In the pilot operationsstripped organic solution (SO) and PLS solution were mixed together inthe first stage of extraction, resulting in a first loaded organicstream and a first raffinate stream. The raffinate from the first stagewas contacted again with a separate SO, resulting in a second loadedorganic and a final raffinate. The two loaded organic streams producedfrom each stage of extraction were then combined as the feed organic toscrubbing; a synthetic (dilute) Co solution, as would be derived from Costripping, was prepared from sulphate salts and used as the feedaqueous. The ratio of Mn in the loaded organic to Co in the aqueous feedsolution was varied from 0.1-36:1.

The objective was to scrub the Mn from the loaded organic prior to thesubsequent Co stripping stages while minimizing the amount of Corecycling to this stage.

Manganese Scrub

The organic and aqueous phases were placed in a one litre stainlesssteel rectangular box immersed in a temperature controlled water bath at30° C., 1:1 O:A and allowed to mix for 180 seconds, while samples weretaken at various intervals.

The results for this test are presented in Table 16 below andgraphically presented in FIG. 10.

TABLE 16 Feed and Product Assays for Mn Scrub Example 7 Co:Mn OrganicAqueous Ratio in Time Co Mn Ni Co Mn Ni Feed s ppm ppm ppm ppm ppm ppm22.6 0 414 70 569 1581 2 8 30 534 5 572 1455 79 21 60 553 0.5 570 144182 21 90 569 0.5 570 1426 85 21 30.4 0 414 70 569 2125 0.5 7 30 541 3577 1967 68 10 60 570 0.5 586 1960 74 9 90 576 0.5 583 1936 72 8 35.8 0414 70 569 2504 0.5 9 30 573 3 604 2526 70 9 60 608 0.5 612 2478 73 1090 656 0.5 641 2462 74 10 1.1 0 418 85 556 96 0.5 531 30 481 17 571 3160 501 60 509 4 597 25 74 505 90 500 <1 587 21 80 525 0.1 0 444 76 61411 0.5 36 30 485 37 687 0.5 38 7 60 454 29 645 0.5 42 4 90 487 30 6880.5 42 3 120 473 28 674 0.5 41 2 150 482 28 692 0.5 45 1 180 464 26 6640.5 44 0.5

Results showing <1 Mn as the aqueous assay are all below the ICPdetection limit so for this purpose we will assume them to be zero. Thetest results indicated that in 90 seconds virtually all of the Mn thatloads onto the organic during processing can be very efficientlyscrubbed off of the organic, with only a minimum quantity of Co requiredin scrub aqueous feed solution. The ratio of Co:Mn in scrub aqueousproduct was low at about 1:4, indicating a minimum use of Co for thisoperation, and hence good possibilities for recycle of this stream.Essentially all of the Mn in LO was replaced by Co, which then wouldproceed to Co stripping, thus making full use of the Co feed toscrubbing.

Conclusions

A small portion of the strip solution produced when selectivelystripping Co from the loaded organic can be efficiently utilized toscrub the Mn from solution allowing a Co rich product to be made withoutthe Mn impurity present. The resulting (Mn-bearing) aqueous solutionfrom this stage can be recycled to the front end of the process, withminimal Co recycle.

Example 8

The example summarizes the testwork in testing the new process.

Nickel, Cobalt and Impurity Extraction

The feed assays for each of the test examples are presented in Table 17.

TABLE 17 Feed Assays for PLS and SO Streams Example 8 Aqueous Organic NiCo Ni Co Feed # ppm ppm ppm ppm 1 20477 1116 0 0 2 21341 1163 0 0 325280 1111 0 0 4 20300 1060 0 0 5 18199 964 0 0 6 19930 1058 — — 7 201991008 0 0 8 25884 1341 0 0 9 22400 1180 — — 10 19320 1020 — — 11 204031039 206 3

Testwork on the new invention was completed in three distinct phases.The first phase consisted of batch testwork where all tests werecompleted with the same procedure as Example 3. The second phaseconsisted of continuous testwork which was operated in either amixer/settler system or pipe reactor system.

Another key operating parameter addressed during the testwork was themethod of neutralization. Neutralization is required during extractionto aintain the pH in the desired range. This must be done by adding aneutralization agent such as caustic.

Three methods of neutralization were tested during the course of thetestwork.

The three methods tested were:

-   -   1. Continuous pH control (pH Control)—caustic was added        continuously to the emulsion in the mixer to maintain a pH value        of 4.5 in both stages of extraction    -   2. Saponification (Sapon.)—a set amount of caustic was added and        mixed with the organic phase prior to contact with the aqueous        phase to maintain a desired pH exiting the pipe reactor    -   3. Partial Saponification and pH control (Partial Sapon.)—a set        amount of caustic (but insufficient), was added to the organic        phase prior to mixing with the aqueous phase and additional        caustic was also added to the mixer to maintain the desired pH        in the settler discharge raffinate.

The continuous pH control method was used throughout a large portion ofthe batch testing and through continuous testing of the mixer/settlersystem. This method was abandoned during continuous phase testing whenit became apparent that a pipe reactor was required to reach the targetNi:Co ratio in the final raffinate.

Table 18 highlights 12 various tests utilizing different equipment andpH control methods.

TABLE 18 Procedures and Assays of Feed and Product Streams for Example 8Operating Parameters E1/E2 Aqueous Organic Mode of Equipment Neut. TimeNi Co Ni Co Test # Feed # Operation Used Method (s) ppm ppm ppm ppm 1 1Batch Batch pH 60 19380 279 978 838 Control 75 17633 28 1620 264 2 2Batch Batch pH 90 20857 428 1800 715 Control 90 18458 57 1374 332 3 3Batch Batch Partial 120 19233 290 773 724 Sapon. 120 16523 24 1270 256 44 Cont. M/S pH 60 — — — — Control 60 22800 169 — — 5 5 Cont. M/S pH 12016874 429 763 674 Control 6 6 Cont. M/S Sapon. 90 18400 366 — — 60 18020174 — — 7 7 Cont. M/S Sapon. 90 17815 526 812 582 8 8 Cont. M/S Partial90 22181 680 607 548 Sapon. 90 20779 316 837 345 9 9 Cont. P/R Sapon. 9020300 198 — — 30 19300 63 — — 10 10 Cont. PM & P/R Sapon. 15 + 90 18630200 1370 740  6 + 30 16680 23 1490 192 11 11 Cont. PM & P/R Sapon. 15 +90 18311 146 2180 945  6 + 75 17259 28 1575 150 Cont. = ContinuousOperations, M/S = Mixer/Settler, P/R = Pipe Reactor and PM = Pre-Mixer

The extraction results and Ni:Co ratios in the final raffinate for eachof the 12 example tests are presented in Table 19.

TABLE 19 Extraction Results for Example 8 Ni:Co Ni Co Test # EquipmentUsed Neut. Method Raffinate % % 1 Batch pH Control 629 12.7 98.7 2 BatchPartial 688 10.4 97.6 Saponification 3 M/S pH Control 59 8.9 69.8 4 M/SpH Control 39 4.2 55.2 5 M/S Saponification 104 12.5 83.6 6 M/SSaponification 43 4.4 45.8 7 M/S Partial 66 5.6 76.4 Saponification 8P/R Saponification 306 23.2 94.7 9 PM & P/R Saponification 83 3.6 78.1713 9.6 89.0 10 PM & P/R Saponification 616 16.7 97.3 M/S =Mixer/Settler, P/R = Pipe Reactor and PM = Pre-Mixer

Test 1 was completed on a batch basis using pH control as theneutralization method. The best results achieved using the pH controlmethod, indicated that a Ni:Co ratio of 629 with 12.7% Ni and 98.7% Coextraction could be achieved. Partial Saponification was then tested asshown in Test No. 2 and the target Ni:Co ratio in the final raffinate(>667:1) was met with less Ni extraction.

Tests 4-7 were all completed on a mixer/settler system in continuousmode. The variable tested during this phase of testwork was the methodof neutralizing the organic. None of these methods of neutralizationproved effective in a mixer/settler system and the end of result of afinal raffinate target of 667:1 Ni:Co was not met.

Since mixer/settlers could not come close to duplicating the resultsachieved during the batch test work, a pipe reactor was tested to tryand achieve a plugged flow system with hopes of duplicating andexceeding the results achieved during the batch testing phase. The firsttest completed which used a pipe reactor is shown in Test No. 8. Thepipe reactor consisted of two pumps and tubing in which if the Reynoldsnumber reached would be 6000. This turbulence should have been enough tomix the two phases however due to the high capacity of the twoperistaltic pumps to obtain the high turbulence the system was notstable, suggesting that it would be difficult to optimize the pipereactor with these pumps. Due to the small tubing used (8 mm innerdiameter) pH control could not be used throughout this test as it wouldhave proved to be too difficult to have multiple caustic addition pointacross the length of the pipe.

Saponification was used as the method of choice for neutralizing theorganic prior to contact with the aqueous solution (PLS) to produce anend raffinate with a pH of 4.5. The results achieved from this testindicated a trend towards those seen during the bench phase with a Ni:Coratio of 306, with high Ni extraction which would need to be overcome.

Tests 9-10 show the use of a pre-mix stage prior to the pipe reactor.This pre-mix stage was added to obtain good mixing and to enable thedispersion to keep throughout the pipe reactor. The use of the pre-mixstage allowed for the dispersion to be consisted throughout the desiredretention time and to achieve the desired Ni:Co in the raffinate.

Conclusions

A Ni:Co ratio of >667:1 can be used with a combination of saponifiedorganic, pre-mix stage and a pipe reactor. This example illustrates howthis invention efficiently separates Co from Ni in Ni sulphide leachliquors, and thus achieves a Ni:Co ratio sufficient for LME GradeNickel.

Example 9

The following example illustrates the continuous results using the newprocess for treating a nickel-cobalt sulphide feed solution. The datafor this test is the averaged data that was taken when operating the newprocess integrated with a hydrometallurgical process for the recovery ofCu and Ni from a bulk sulphide concentrate over a one month period.

Cobalt Extraction

The extraction system consisted of two trains in a series-parallelconfiguration where the PLS flowed through each of the two mixers/pipereactors/settlers in series, and the organic flowed through in parallel,both at 439 mL/min. That is, freshly saponified organic entered each ofthe two trains independently, and after a single pass through anextraction stage, recombined in the Loaded Organic Tank. Due to thesmall scale used during this phase of testing, pre-mix boxes were usedprior to the in-line mixer/pipe reactor to increase the aqueous/organicdispersion in the tubing. E1 pre-mix box had a retention time of 15seconds and a pipe reactor/inline mixer (pipe reactor consisted of 8 mmTygon tubing with a PVC inline mixer of ½″ diameter×11″ long) with atotal retention time of 105 seconds, whereas the E2 mixer had aretention time of 6 seconds and a pipe reactor/inline mixer retention of105 seconds. The raffinate produced from the second stage has a Ni:Coratio ≧667:1 which is sent to the Nickel Hydroxide Precipitationcircuit. The loaded organics from each stage are combined and sentthrough the strip circuit. The pH of the extraction circuit wasmonitored regularly, and controlled in the target range of 3.9-4.1 forthe first stage and 4.1-4.3 for the second stage by caustic addition tothe stripped organic utilizing 30% NaOH solution making saponifiedorganic. The operating parameters for the extraction stages arepresented in Table 20.

TABLE 20 Operating Parameters for Example 9 Parameter Units E1 E2 Temp °C. 30 30 External O:A — 1:1 1:1 Mixer O:A — 1:1 1:1 Mixer PC — A/C A/CMixer RT S 15 6 In-line Mixer RT S 105 105 Settler RT Min 9.8 9.8 PC =phase continuity, RT = retention time, A/C = aqueous continuous, O/C = organic continuous.

The feed and product solutions for the series parallel extraction stagesare presented in Table 21 and Table 22.

TABLE 21 Feed Assays for Example 9 Aqueous Organic Ni Ca Cl Co Mg Mn NaNi Ca Co Mg Mn Na Stream ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppmppm PLS/SO 21211 602 7918 780 109 4128 897 189 1 3 <1 <1 1186

TABLE 22 Product Assays for Example 9 Aqueous Organic Ni Co Na Ni Ca CoMn Mg Na Stage ppm ppm ppm ppm ppm ppm ppm ppm ppm 1 19938 222 2631 13941 589 3 <1 <1 2 17658 24 4506 2150 1 222 3 <1 <1

The extraction results achieved for continuous operations are presentedin Table 23.

TABLE 23 Extraction Results Example 9 Ni Ca Co Mg Mn g NaOH/g Co Stage %% % % % extracted 1 5.7 1.2 71.5 0 4.6 8.3 2 9.8 0.2 89.2 0 0 Overall15.5 1.3 96.9 0 1.8

Conclusions

The new invention successfully extracted Co in two stages and produced araffinate with a Ni:Co ratio of 735:1 with minimal co-extraction ofimpurities.

Example 10

The following example illustrates the use of the new process forextracting Co from a nickel-cobalt laterite feed solution in continuousmode.

Nickel, Cobalt and Impurity Extraction

The same series parallel extraction was used for this example that wasused for example 8, with sulphide leach liquor. However, the retentiontimes used were slightly different with the first stage of extractionhad a 6 seconds pre-mix time and 114 seconds retention time in theinline mixer and the second stage of extraction had 15 seconds pre-mixtime and 136 seconds in the inline mixer. The feed and product solutionsfor the series parallel extraction stages are presented in Table 24 andTable 25.

TABLE 24 Feed Assays for Continuous Laterite Feed Operations Example 10Aqueous Organic Ni Ca Cl Co Mg Mn Na Ni Ca Co Mg Mn Na Stream ppm ppmppm ppm ppm ppm ppm ppm ppm ppm ppm ppm ppm PLS/SO 7220 202 2632 7801591 966 674 148 1 3 <1 <1 <1

TABLE 25 Product Assays for Continuous Laterite Feed Operations Example10 Aqueous Organic Ni Co Na Ni Ca Co Mn Mg Na Stage ppm ppm ppm ppm ppmppm ppm ppm ppm 1 6870 110 1585 611 1 613 58 1 107 2 6103 6 2216 798 <1109 135 1 105

The extraction results achieved during these tests are presented inTable 26.

TABLE 26 Extraction Results for Continuous Laterite Feed OperationsExample10 g NaOH/g Ni Ca Co Mg Mn Co Stage % % % % % extracted E1 6.40.0 84.1 0.0 0.6 3.4 E2 9.5 0.0 94.8 0.0 3.6 Overall 15.9 0.0 99.1 0.02.2

Conclusions

The new process successfully removed >99.1% of Co in two stages ofextraction, and thus produced a raffinate with a Ni:Co ratio of 1017:1.There was a minor amount of extraction of Mn which can be readilyscrubbed from the organic as shown in Example 7. The caustic consumptionof 3.4 g NaOH per g Co is very low, and partly reflects the lower Ni:Coratio in the PLS, compared to Example 9.

Example 11

This example illustrates another embodiment of the invention, i.e.selectively stripping Co from the loaded organic stream (containingNi+Co), that was produced in the extraction stages.

Cobalt Selective Stripping

The fresh organic solution containing 4.7 v/v % Versatic 10 and 14 v/v %LIX 63 in Shellsol D70 was contacted with fresh aqueous solution (PLS)at pH 4.5, 30° C. and on O:A ratio of 1:1 for 60 seconds resulting inthe loaded organic solution for this test.

Two separate stripping tests were carried out, with varying feed acidinto the strip. The loaded organic solution was stripped with aqueoussolutions containing either 3 or 5 g/L H₂SO₄ at an O:A ratio of 1:1 and30° C. for 180 seconds; samples were taken at intervals to examine thestripping kinetics. The strip solutions also contained 0.2M NaS₂O₄ so asto increase the ionic strength of the strip solution, and thus improvephase separation. The results of these tests are shown in Table 27 andFIG. 11.

The objective of this test was to produce a strip product solution witha Co:Ni ratio of 1:1 or greater.

TABLE 27 Co and Ni Stripping Kinetics with 3 and 5 g/L Sulphuric AcidExample 11 Organic Strip Liquor Stripping Acid Time Co Ni Co Ni Co Nig/L s (g/L) (g/L) (g/L) (g/L) Co:Ni (%) (%) 3 0 0.739 0.975 0 0 — — — 150.397 0.914 0.342 0.061 5.6 46.3 6.27 30 0.300 0.937 0.439 0.038 11.659.4 3.86 45 0.208 0.942 0.531 0.033 16.1 71.9 3.36 60 0.189 0.933 0.5500.042 13.1 74.4 4.30 90 0.140 0.892 0.599 0.083 7.2 81.0 8.48 120 0.0960.862 0.643 0.113 5.7 87.0 11.5 180 0.051 0.858 0.688 0.117 5.9 93.112.0 5 0 0.711 0.756 0 0 — — — 15 0.088 0.607 0.603 0.159 3.79 89.6 19.730 0.080 0.533 0.608 0.127 4.79 88.7 29.4 45 0.038 0.441 0.630 0.1464.32 94.6 41.6 60 0.041 0.478 0.642 0.148 4.34 94.3 36.7 90 0.025 0.4870.675 0.138 4.89 96.4 35.6 180 0.006 0.447 0.696 0.166 4.19 99.1 40.9

At 5 g/L sulphuric acid at 1:1 and 30° C., nearly 90% Co was stripped in15 seconds and 95% in 60 seconds, indicating very fast Co strippingkinetics. The Ni stripping kinetics was much slower with only 40% beingstripped in 60 seconds and then the stripping efficiency leveled off.

At 3 g/L sulphuric acid, the Co stripping kinetics was slower than with5 g/L sulphuric acid, but it was more selective. In 60 seconds, 74% Coand only 4% Ni were stripped, leading to a Co:Ni ratio in aqueousproduct of 13:1. At 3 minutes, over 93% Co and 12% Ni were stripped,with a product of 6:1 ratio. This indicates that selective stripping ofCo over Ni was much more effective using lower acidity strip solutionthan higher acidity strip solution.

Conclusions

The majority of Co can be easily and selectively stripped fromNi-bearing loaded organic at low acid concentrations, low temperaturesand short retention times.

Example 12

This example illustrates the effect of temperature on the strippingkinetics for stripping Ni from the Co-depleted organic produced in theCo selective stripping stage.

Nickel Stripping

The loaded organic solution containing 4.7 v/v % Versatic 10 and 14 v/v% LIX 63 in Shellsol D70 was first selectively stripped of the majorityof the Co prior to this example. The Co-depleted loaded organic solutionwas stripped with 50 g/L H₂SO₄ at an O:A ratio of 1:1 and temperaturesof 30, 40, 45 and 50° C. The stripping kinetics for these tests is shownin Table 28 and FIG. 12.

TABLE 28 Ni Stripping Kinetics at Varying Temperatures Example 1216Loaded Organic Stripped Organic Stripping Temperature Time Co Ni Co NiCo Ni ° C. s (g/L) (g/L) (g/L) (g/L) (%) (%) 30 0 — 1.212 — — — — 30 —1.212 — — — 20.7 60 — 1.212 — — — 27.6 120 — 1.212 — — — 29.7 180 —1.212 — — — — 240 — 1.212 — — — 33.8 300 — 1.212 — — — — 480 — 1.212 — —— 52.4 40 0 — 1.212 — — — — 30 — 1.212 — 0.862 — 17.4 60 — 1.212 — 0.810— 32.2 120 — 1.212 — 0.589 — 47.9 180 — 1.212 — 0.341 — 64.5 240 — 1.212— — — — 300 — 1.212 — 0.001 — 86.8 480 — 1.212 — 0.001 — 94.2 45 0 0.1251.175 — — — — 30 0.125 1.175 0.022 0.674 82.1 42.6 60 0.125 1.175 0.0200.609 83.8 48.2 120 0.125 1.175 0.017 0.435 86.2 63.0 180 0.125 1.1750.015 0.228 88.2 80.6 240 0.125 1.175 — — — — 300 0.125 1.175 0.0120.035 90.8 97.0 480 0.125 1.175 0.009 0.005 93.2 99.6 50 0 0.123 1.210 —— — — 30 0.123 1.210 0.013 0.633 89.6 47.5 60 0.123 1.210 0.009 0.45592.4 62.1 120 0.123 1.210 0.007 0.183 94.1 84.1 180 0.123 1.210 0.0060.053 95.3 95.7 240 0.123 1.210 — — — — 300 0.123 1.210 0.005 0.006 96.099.6 480 0.123 1.210 0.005 0.007 96.2 99.6

This test indicates the Ni stripping efficiency is very temperaturedependant. After 5 minutes of stripping, the Ni stripping efficiencyincreased from 86.8% at 40° C. to 97% at 45° C. and further to 99.6% at50° C.

Conclusions

The Ni can be stripped down to very low levels at high temperatures. Thehalf life of LIX 63 at 45° C. and 50° C. under these strip conditionscould be over 1.5 years and over 1 year respectively. Therefore,stripping at higher temperature can be an option for complete Nistripping.

Example 13

This example illustrates the use of selectively stripping Co from theloaded organic in continuous mode, with varying acid concentration andvolume.

Cobalt Selective Stripping

The loaded organic solution contained 3.8 v/v % Versatic 10 and 19.5 v/v% LIX 63 in Shellsol D70 and was produced from continuous extractionoperations. The loaded organic solution was stripped with acid at 30-34°C. in two different arrangements:

-   -   a) With a relatively large volume of dilute strip solution or    -   b) With a small volume of more concentrate acid strip solution.

The results from these tests are presented in Table 29. The objectivesincluded producing a strip product solution with a Co:Ni ratio of 1:1 orgreater, maximize Co stripping, with a maximum [Co] and minimum acidity(for greater ease of Co recovery downstream).

TABLE 29 Co and Ni Stripping Results with 10-40 g/L Sulphuric AcidExample 13 Organic Feed Strip Liquor Stripping Test O:A Temp Time AcidCo Ni Co Ni Co Ni No Ratio ° C. s g/L (g/L) (g/L) (g/L) (g/L) Co:Ni (%)(%) 1  5:1 30 4 10 0.669 1.49 0.606 0.154 3.9 94.6 18.5 2  5:1 32 4 100.466 1.28 1.37 0.709 1.9 92.0 16.3 3  5:1 33 4 10 0.488 1.62 2.18 0.7602.9 86.4 11.5 4  5:1 34 5 10 0.558 1.67 2.22 0.876 2.5 87.2 10.2 5  5:134 5 10 0.516 1.38 2.15 1.03 2.1 85.3 2.3 6 20:1 34 4 40 0.515 1.65 5.552.82 2.0 85.1 9.8 7 24:1 30 5 40 0.406 1.66 7.12 3.25 2.2 74.2 0.0 825:1 31 4 40 0.494 1.66 7.72 3.11 2.5 82.6 2.4 9 20:1 32 4 40 0.533 1.877.35 3.46 2.1 85.0 13.0 10 20:1 31 4 40 0.495 1.36 6.91 2.58 2.7 83.86.9

Co strip efficiency was in the range of 85-94% with a strip O:A ratio of5:1 using a strip solution containing 10 g/L sulphuric acid. The Costrip efficiency decreased to a range of 82-85% in most cases with astrip O:A ratio of 20:1 using a strip solution containing 40 g/Lsulphuric acid. The Co:Ni ratio increased from a range of 0.25-0.36 inthe loaded organic to 2.0-2.9 in the strip liquor in most cases. Thismeans that after the Co selective stripping the target of a Co:Ni ratioof 1 in the strip liquor was easily achieved.

Conclusions

The majority of Co was selectively stripped from Ni at low acidconcentrations and O:A ratios of 5:1. It also suggests that it might bepossible to further limit the Ni stripping (in this Co Strip) by runninga two stage counter-current strip by running the conditions set out inTest Number 7 which indicated no co-stripping of the Ni with the Co.

Example 14

This example illustrates the continuous mode of stripping Ni from theCo-depleted organic produced from the Co selective stripping stage.

Ni Stripping

The Co-depleted organic used for this test came from the Co selectivestripping stage as described in the previous example. As in the previousexample, the volume and acidity of the strip solution was varied.

The Co-depleted organic solution was stripped with a strip solutionscontaining 10-80 g/L H₂SO₄ at an O:A ratio of 3-29:1 and 45-53° C. Theresults from these tests are presented in Table 30. The objective ofthis test was to produce a stripped organic solution with a residual Niconcentration of <0.25 g/L to limit the effects on the subsequentextraction stages, as well as attempting to make a Ni strip solutionwith a maximum [Ni] in minimum volume, etc as before for Co.

The objective of this test was to produce a strip organic solution witha residual Ni concentration of <0.25 g/L to limit the effects on thesubsequent extraction stages, as well as attempting to make a Ni stripsolution with maximum [Ni] in minimum volume, etc as before for Co.

TABLE 30 Co and Ni Stripping Results with 10-80 g/L Sulphuric AcidExample 14 Organic Organic Strip In Out Liquor Stripping Test O:A TempTime Acid Co Ni Co Ni Co Ni Co Ni No Ratio ° C. s g/L (g/L) (g/L) (g/L)(g/L) (g/L) (g/L) (%) (%) 1  5:1 45 8 10 0.038 1.07 0.002 0.139 0.1942.03 95.4 87.2 2  3:1 48 8 10 0.067 1.44 0.000 0.490 0.428 3.54 99.665.9 3  5:1 49 10 10 0.071 1.50 0.000 0.529 0.359 3.98 99.7 54.7 4 20:153 10 40 0.076 1.35 0.000 0.276 0.445 3.70 99.7 79.5 5 20:1 48 8 400.077 1.49 0.000 0.437 1.04 7.59 99.7 40.6 6 29:1 50 8 60 0.105 1.710.000 0.535 2.41 16.8 99.8 68.6 7 26:1 49 8 80 0.086 1.62 0.003 0.3362.13 15.8 99.7 79.3 8 20:1 50 8 80 0.080 1.63 0.000 0.193 2.43 24.3 99.788.2 9 20:1 51 8 45 0.080 1.27 0.002 0.226 2.54 19.0 98.1 82.2

As indicated in Table 30 a 20:1 O:A ratio and 45 g/L sulphuric acid wasable to achieve the desired Ni stripped organic tenors of less than 0.25g/L.

Conclusions

Stripping the co-depleted organic solution to less than 0.25 g/L Ni canbe achieved, to produce a concentrated Ni solution of about 20 g/l Ni,with a pH value of <1.

Example 15

This example illustrates the use of bulk strip with a final Co/Niproduct solution.

Cobalt and Nickel Stripping

The loaded organic solution contained 3.8 v/v % Versatic 10 and 19.5 v/v% LIX 63 in Shellsol D70 and was produced from continuous extractionoperations. The loaded organic solution was stripped at two differentO:A ratios while maintaining the strip solution at 80 g/L H₂SO₄ in athree stage counter current operation, as shown in FIG. 2. The resultsfrom these tests are presented in Table 31.

The objective of these tests was to produce a final Co PE possessing a<4:1 Ni:Co ratio while ensuring a Co PE free acid content of <1.0 g/L.It was also imperative that the final stripped organic heading back toextraction had a nickel tenor in the range of 180-250 ppm for thepurposes of optimal extraction.

TABLE 31 Continuous Bulk Strip Results Example 15 Feed Feed Prod. StripFree Organic Aqueous Organic Prod. Aq Efficiency O:A Acid (ppm) (ppm)(ppm) (ppm) (%) Ratio Stage (g/L) Ni Co Ni Co Ni Co Ni Co Ni Co 20 Acid81.1 — — — — — — — — — — Feed S1 0.38 1880 497 19480 3770 1920 334 166706090 −1.4 32.8 S2 2.7 1920 334 9880 277 1250 26 19480 3770 27.8 92.2 S330.1 1250 26 0 0 253 7 9880 277 79.8 73.1 Overall — 1880 497 0 0 253 716670 6090 86.5 98.6 15 Acid 81.1 — — — — — — — — — — Fd S1 0.25 1940534 19000 3660 1990 384 16800 6470 −1.6 28.1 S2 5.07 1990 384 12200 3401130 33 19000 3660 30.7 91.4 S3 35.6 1130 33 0 0 178 3 12200 340 84.390.9 Overall — 1940 534 0 0 178 3 16800 6470 90.8 99.4

When utilizing an O:A ratio of 20:1 and a free acid concentration of81.1 g/L in a three stage counter current operation 99.5% of the acidfeed into the circuit was utilized leaving 0.25 g/L FA in the Co/Niproduct stream. This resulted in an overall stripping efficiency of90.8% Ni and 99.4% Co. This operation allowed the Ni in the organic tobe slightly less than the target of 180-250 ppm which indicates that anO:A ratio between 15 and 20 would sufficiently meet the required targetsfor this circuit.

Conclusions

A non selective stripping approach can be taken to successfully stripthe Co and Ni from the organic solution while producing a product <3:1Ni:Co.

Example 16

The following example illustrates the continuous results using the newinvention for treating a nickel-cobalt sulphide feed solution. The datafor this example is the averaged data that was taken when operating thenew invention integrated with the CESL Cu/Ni Process over a one monthperiod.

Co Extraction and Bulk Stripping

The extraction system was operated under the same conditions aspresented in Example 8 and the strip circuit was operated in a threestage counter current mode with two acid feeds. A flow diagram for thiscircuit is presented in FIG. 13.

The operating parameters for the extraction and stripping stages arepresented in Table 32.

TABLE 32 Operating Parameters for Example 16 Parameter Units E1 E2 S1 S2S3 Temp ° C. 30 30 30 50 50 External O:A — 1:1 1:1 17:1 17:1 17:1 MixerO:A — 1:1 1:1  1:1  1:1  1:1 Mixer PC — A/C A/C A/C A/C A/C Mixer RT s15 6 5 min 10 min 10 min In-line Mixer RT s 105 105 — — — Settler RT min9.8 9.8 19.8 20.3 20.3 PC = phase continuity, RT = retention time, A/C =aqueous continuous, O/C = organic continuous

The feed and product solutions for the series parallel extraction stagesare presented in Table 33 and Table 34.

TABLE 33 Extraction Feed Assays for Example 16 Aqueous Organic Ni Ca ClCo Mg Mn Na Ni Ca Co Mg Mn Na Stream ppm ppm ppm ppm ppm ppm ppm ppm ppmppm ppm ppm ppm PLS/SO 21211 602 7918 780 109 4128 897 189 1 3 <1 <11186

TABLE 34 Product Assays for Example 16 Aqueous Organic Ni Co Na Ni Ca CoMn Mg Na Stage ppm ppm ppm ppm ppm ppm ppm ppm ppm 1 19938 222 2631 13941 589 3 <1 <1 2 17658 24 4506 2150 1 222 3 <1 <1

The extraction results achieved for continuous operations are presentedin Table 35.

TABLE 35 Extraction Results for Example 16 g NaOH/g Ni Ca Co Mg Mn CoStage % % % % % extracted 1 5.7 1.2 71.5 0 4.6 8.3 2 9.8 0.2 89.2 0 0Overall 15.5 1.3 96.9 0 1.8

The strip circuit employed a two acid feed system to enable the tenorsin the stripped organic to meet the target of 200 ppm Ni and with thethought that the strip product solution from S3 could be integrated withthe final raffinate solution from the extraction stages which wouldproduce a Ni:Co ratio >667:1.

The strip flowsheet is presented in FIG. 13.

FIG. 13 is a simplified flow diagram illustrating another bulk strippingoption with two acid feed streams. The stripping comprises three stages130, 132 and 134. The combined loaded organic 128 from extraction isintroduced into the first stage stripping 130 and then passes to stages132 and 134 to exit as stripped organic 146 from the third stage 134.

In the third stage of stripping 134 a strong acid feed solution 136 withfree acid content of 37:3 g/L is added to remove the remaining Ni formthe partially stripped organic. The product aqueous stream from thisstage 134 is split, as shown at 138, with a 50% portion being a Niproduct solution 140 with low levels of Co and the remaining stream isadded to the second stage stripping 132.

In the second stage 132 another dilute acid stream 142 is introduced toenable the maintenance of a 20:1 organic:aqueous ratio across thecircuit.

The first stage of stripping 130 produced the Co rich product 144 with aNi:Co ratio of about 3:1. To enable the product stream 140 produced in134 to be integrated with the final raffinate product in E2 the cobaltcontent could not be higher than 4 ppm Co. Average feed productcompositions for the organic and aqueous streams for the strippingstages are presented in Table 36.

TABLE 36 Average Feed and Product Compositions for Strip Circuit Example16 FA Cl Mn Mg Co Ni Stream g/L ppm ppm ppm ppm ppm LO —  3 2 BDL 4041733 Co Rich  1.1 533 2 88 5185 16975  Product S2 Acid Feed 48.3 — — — —— SO — — <1  <1 2  183 S3 Acid Feed 37.3 — — — — — S3 Product 36.6 116 2 4 143 8512

The average cobalt assay for the S3 product solution 140 was much higherthan the desired 4 ppm at 143 ppm. This product would not be able to beintroduced to the E2 raffinate because the Ni:Co ratio would be reducedfrom 736 to 475.

The performances of the three stage counter-current strip are outlinedin Table 37.

As noted in the table the overall stripping efficiency of 89.4% Ni and99.5% Co realized Ni tenors in the resulting stripped organic to be <200ppm and the Co rich product had a Ni:Co ratio of 3:1.

TABLE 37 Average Stripping Efficiencies for Example 16 Feed Feed Prod.Strip Organic Aqueous Organic Prod. Aq Efficiency (ppm) (ppm) (ppm)(ppm) (%) Stage Ni Co Ni Co Ni Co Ni Co Ni Co S1 1733 415 17580 18781724 139 16975 5185 0.30 67.0 S2 1724 139 4256 72 746 9 17580 1878 47.993.5 S3 746 9 0 0 183 2 8512 143 75.5 77.8 Overall 1733 415 0 0 183 2 —— 89.4 99.5

Conclusions

The new invention successfully extracted Co in two stages and produced araffinate with a Ni:Co ratio of 735:1 with minimal co-extraction ofimpurities and stripped the organic solution to <200 ppm Ni. The productsolution produced from S3 had too much Co in solution to integrate withthe final raffinate product. A Bulk strip where one acid feed is usedproducing one product solution would need to be used for this flowsheetas documented in Example 14.

TABLE 38 Effect of temperature on Co and Ni extraction at pH 4.5 and anA/O ratio of 1:1 Temp. Con. Concentrations (g/L) (° C.) (g/L) 15 (s) 30(s) 45 (s) 60 (s) 90 (s) 180 (s) 22 Co 0.40 0.57 0.69 0.76 0.86 0.94 Ni0.24 0.40 0.54 0.67 0.86 1.42 Co/Ni ratio 1.67 1.43 1.28 1.13 1.00 0.6630 Co 0.25 0.43 0.58 0.69 0.79 0.56 Ni 0.24 0.38 0.55 0.69 1.02 1.55Co/Ni ratio 1.07 1.13 1.05 0.99 0.77 0.36 35 Co 0.44 0.63 0.76 0.84 0.810.13 Ni 0.36 0.61 0.85 1.14 1.47 1.94 Co/Ni ratio 1.22 1.03 0.89 0.740.55 0.07 40 Co 0.54 0.79 0.59 0.41 0.09 — Ni 0.53 1.13 1.47 1.63 1.84 —Co/Ni ratio 1.02 0.70 0.40 0.25 0.05 —

The above-described embodiments of the invention are intended to beexamples only. Alterations, modifications and variations can be effectedto the particular embodiments by those of skill in the art withoutdeparting from the scope of the invention, which is defined solely bythe claims appended hereto.

What is claimed is:
 1. A process for separating and recovering Co fromNi in an aqueous feed solution comprising sulphates, comprising:subjecting the feed solution comprising the sulphates and containing Niand Co and impurities comprising Mg and Mn to extraction with anextractant and using kinetic differences between Ni and Co in theextraction by controlling duration of the extraction, thereby separatingthe Co from the Ni, and from the Mg and the Mn, whereby a major portionof Co and a minor portion of Ni is extracted from the feed solution toproduce a loaded extractant and a Co-depleted raffinate containing Niwherein the loaded extractant is enriched in Co and depleted in Nicompared to the feed solution, wherein the concentration of Ni in thefeed solution is greater than the concentration of Co, and wherein theduration of extraction is about 30 to 240 seconds, wherein the molarratio of (Mg+Mn):Co in the feed solution is at least 6:1.
 2. The processof claim 1, wherein the feed solution is acidic.
 3. The process of claim1, wherein the feed solution is derived from leaching of ores orconcentrates containing Ni and Co.
 4. The process of claim 1, whereinthe duration of the extraction is controlled by limiting the duration ofthe extraction to a period wherein a maximum ratio of Co:Ni has beenextracted.
 5. The process of claim 1, wherein the duration of theextraction is controlled by limiting the duration of the extraction to aperiod wherein a maximum ratio of Ni:Co remains unextracted.
 6. Theprocess of claim 1, comprising a plurality of extractions, wherein theduration of each of the extractions is about 30 to 240 seconds.
 7. Theprocess of claim 1, wherein the extraction is carried out at atemperature of about 15° C. to 50° C.
 8. The process of claim 1, whereinthe extraction is carried out at an organic to aqueous ratio of about0.5:1 to 4:1.
 9. The process of claim 1, further comprising controllingpH of the extraction by adding an alkali.
 10. The process of claim 9,wherein the alkali is added to the extractant in an amount of about 0.01to 0.3 grammole per litre.
 11. The process of claim 1, wherein theextraction is carried out at a pH of about 2.5 to 5.5.
 12. The processof claim 1, further comprising subjecting the extractant tosaponification prior to contact with the aqueous feed solution, whereinthe extractant is contacted with an alkali to pre-neutralize or saponifythe extractant.
 13. The process of claim 1, wherein the extractantcomprises a mixture of at least two extractants and one of the twoextractants is a carboxylic acid and the other of the two extractants isa hydroxyoxime.
 14. The process of claim 13, wherein the carboxylic acidcomprises 2-methyl, 2-ethyl heptanoic acid and the carboxylic acidextractant percentage is about 2 to 20 v/v % with reference to totalvolume of the extractants and a diluent.
 15. The process of claim 14,wherein the hydroxyoxime comprises 5,8-diethyl-7-hydroxy-6-dodecanoneoxime and the hydroxyoxime extractant percentage is about 4 to 40 v/v %with reference to total volume of the extractants and a diluent.
 16. Theprocess of claim 1, wherein the extraction is carried out in two or moresuccessive stages and raffinate from the first stage is contacted byfresh or recycled stripped extractant in the second stage to extractmore Co, thereby to further improve the Ni:Co ratio in raffinateproduced by the second stage.
 17. The process of claim 1, furthercomprising stripping the Co and Ni from the loaded extractant with anacidic strip solution having a pH of about 1.5 to about 2.5 to produce aCo and Ni product solution and a stripped extractant which is recycledto the extraction, wherein the stripping is carried out a temperature ofabout 30° C. to 60° C., in 2 to 4 stages, and for a duration of about 3to 15 minutes per stage.
 18. The process of claim 1, wherein the loadedextractant is subjected to stripping with an acidic solution to producean aqueous product solution containing Ni and Co and further comprisingsubjecting the aqueous product solution to a further extraction stage toproduce a Co-loaded extractant, which is subjected to further strippingto produce a Co product solution, and a second Ni raffinate which iscombined with the Co-depleted raffinate from the Ni and Co extraction toproduce a Ni product solution.
 19. The process of claim 1, furthercomprising selectively stripping Co from the loaded extractant with adilute acidic strip solution to produce a Co solution having a pH ofabout 1 to 2.5 and a partially stripped extractant and then subjectingthe Co solution to a second Co extraction to produce a Co-loadedextractant and a second Ni raffinate, wherein the selective stripping iscarried out a temperature of about 20 to 40° C., in 1 or 2 stages, andfor a duration of about 1 to 10 minutes per stage.
 20. The process ofclaim 19, further comprising stripping Ni from the partially strippedextractant with a stronger acidic strip solution to produce a Nisolution having a pH of about 1.0 to 2.0 and a stripped extractant andrecycling the stripped extractant to the extraction, wherein thestripping is carried out a temperature of about 30° C. to 60° C., in 1to 6 stages, and for a duration of about 3 to 15 minutes per stage. 21.The process of claim 20 further comprising subjecting the Ni-solution toa third Co-extraction with stripped extractant from the secondCo-extraction to produce a further Co-loaded extractant and a third Niraffinate which is combined with both the Co-depleted raffinate from theextraction and the second Ni raffinate to produce a Ni product solution.22. The process of claim 1, further comprising scrubbing any Mn presentfrom the loaded extractant with a scrub aqueous solution comprising partof the solution enriched in Co or a part of the raffinate, the scrubaqueous solution having a cobalt to manganese ratio of about 10:1 to0.75:1, wherein the scrubbing is carried out at a temperature of about20° C. to 40° C.
 23. The process of claim 1, wherein the extraction iscarried out using an inline mixer or pipe mixer for mixing the feedsolution and the extractant.
 24. A hydrometallurgical process for therecovery of Ni or Co or both Ni and Co from a feed material containingNi and Co, comprising subjecting the feed material to acid leaching toobtain a resultant acid leach solution and treating the acid leachsolution according to the process of claim 1.